Optimized liquid-phase oxidation

ABSTRACT

Disclosed is an optimized process and apparatus for more efficiently and economically carrying out the liquid-phase oxidation of an oxidizable compound. Such liquid-phase oxidation is carried out in a bubble column reactor that provides for a highly efficient reaction at relatively low temperatures. When the oxidized compound is para-xylene and the product from the oxidation reaction is crude terephthalic acid (CTA), such CTA product can be purified and separated by more economical techniques than could be employed if the CTA were formed by a conventional high-temperature oxidation process.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims priority to U.S. Provisional Application Ser.No. 60/606,732, filed Sep. 2, 2004 and 60/631,519, filed Nov. 29, 2004,the disclosures of which are incorporated herein by reference in theirentirety to the extent they do not contradict statements herein.

FIELD OF THE INVENTION

This invention relates generally to a process for the liquid-phase,catalytic oxidation of an aromatic compound. One aspect of the inventionconcerns the partial oxidation of a dialkyl aromatic compound (e.g.,para-xylene) to produce a crude aromatic dicarboxylic acid (e.g., crudeterephthalic acid), which can thereafter be subjected to purificationand separation. Another aspect of the invention concerns an improvedbubble column reactor that provides for a more effective and economicalliquid-phase oxidation process.

BACKGROUND OF THE INVENTION

Liquid-phase oxidation reactions are employed in a variety of existingcommercial processes. For example, liquid-phase oxidation is currentlyused for the oxidation of aldehydes to acids (e.g., propionaldehyde topropionic acid), the oxidation of cyclohexane to adipic acid, and theoxidation of alkyl aromatics to alcohols, acids, or diacids. Aparticularly significant commercial oxidation process in the lattercategory (oxidation of alkyl aromatics) is the liquid-phase catalyticpartial oxidation of para-xylene to terephthalic acid. Terephthalic acidis an important compound with a variety of applications. The primary useof terephthalic acid is as a feedstock in the production of polyethyleneterephthalate (PET). PET is a well-known plastic used in greatquantities around the world to make products such as bottles, fibers,and packaging.

In a typical liquid-phase oxidation process, including partial oxidationof para-xylene to terephthalic acid, a liquid-phase feed stream and agas-phase oxidant stream are introduced into a reactor and form amulti-phase reaction medium in the reactor. The liquid-phase feed streamintroduced into the reactor contains at least one oxidizable organiccompound (e.g., para-xylene), while the gas-phase oxidant streamcontains molecular oxygen. At least a portion of the molecular oxygenintroduced into the reactor as a gas dissolves into the liquid phase ofthe reaction medium to provide oxygen availability for the liquid-phasereaction. If the liquid phase of the multi-phase reaction mediumcontains an insufficient concentration of molecular oxygen (i.e., ifcertain portions of the reaction medium are “oxygen-starved”),undesirable side-reactions can generate impurities and/or the intendedreactions can be retarded in rate. If the liquid phase of the reactionmedium contains too little of the oxidizable compound, the rate ofreaction may be undesirably slow. Further, if the liquid phase of thereaction medium contains an excess concentration of the oxidizablecompound, additional undesirable side-reactions can generate impurities.

Conventional liquid-phase oxidation reactors are equipped with agitationmeans for mixing the multi-phase reaction medium contained therein.Agitation of the reaction medium is supplied in an effort to promotedissolution of molecular oxygen into the liquid phase of the reactionmedium, maintain relatively uniform concentrations of dissolved oxygenin the liquid phase of the reaction medium, and maintain relativelyuniform concentrations of the oxidizable organic compound in the liquidphase of the reaction medium.

Agitation of the reaction medium undergoing liquid-phase oxidation isfrequently provided by mechanical agitation means in vessels such as,for example, continuous stirred tank reactors (CSTRs). Although CSTRscan provide thorough mixing of the reaction medium, CSTRs have a numberof drawbacks. For example, CSTRs have a relatively high capital cost dueto their requirement for expensive motors, fluid-sealed bearings anddrive shafts, and/or complex stirring mechanisms. Further, the rotatingand/or oscillating mechanical components of conventional CSTRs requireregular maintenance. The labor and shutdown time associated with suchmaintenance adds to the operating cost of CSTRs. However, even withregular maintenance, the mechanical agitation systems employed in CSTRsare prone to mechanical failure and may require replacement overrelatively short periods of time.

Bubble column reactors provide an attractive alternative to CSTRs andother mechanically agitated oxidation reactors. Bubble column reactorsprovide agitation of the reaction medium without requiring expensive andunreliable mechanical equipment. Bubble column reactors typicallyinclude an elongated upright reaction zone within which the reactionmedium is contained. Agitation of the reaction medium in the reactionzone is provided primarily by the natural buoyancy of gas bubbles risingthrough the liquid phase of the reaction medium. This natural-buoyancyagitation provided in bubble column reactors reduces capital andmaintenance costs relative to mechanically agitated reactors. Further,the substantial absence of moving mechanical parts associated withbubble column reactors provides an oxidation system that is less proneto mechanical failure than mechanically agitated reactors.

When liquid-phase partial oxidation of para-xylene is carried out in aconventional oxidation reactor (CSTR or bubble column), the productwithdrawn from the reactor is typically a slurry comprising crudeterephthalic acid (CTA) and a mother liquor. CTA contains relativelyhigh levels of impurities (e.g., 4-carboxybenzaldehyde, para-toluicacid, fluorenones, and other color bodies) that render it unsuitable asa feedstock for the production of PET. Thus, the CTA produced inconventional oxidation reactors is typically subjected to a purificationprocess that converts the CTA into purified terephthalic acid (PTA)suitable for making PET.

One typical purification process for converting CTA to PTA includes thefollowing steps: (1) replacing the mother liquor of the CTA-containingslurry with water, (2) heating the CTA/water slurry to dissolve the CTAin water, (3) catalytically hydrogenating the CTA/water solution toconvert impurities to more desirable and/or easily-separable compounds,(4) precipitating the resulting PTA from the hydrogenated solution viamultiple crystallization steps, and (5) separating the crystallized PTAfrom the remaining liquids. Although effective, this type ofconventional purification process can be very expensive. Individualfactors contributing to the high cost of conventional CTA purificationmethods include, for example, the heat energy required to promotedissolution of the CTA in water, the catalyst required forhydrogenation, the hydrogen stream required for hydrogenation, the yieldloss caused by hydrogenation of some terephthalic acid, and the multiplevessels required for multi-step crystallization. Thus, it would bedesirable to provide a CTA product that could be purified withoutrequiring heat-promoted dissolution in water, hydrogenation, and/ormulti-step crystallization.

OBJECTS OF THE INVENTION

It is, therefore, an object of the present invention to provide a moreeffective and economical liquid-phase oxidation reactor and process.

Another object of the invention is to provide a more effective andeconomical reactor and process for the liquid-phase catalytic partialoxidation of para-xylene to terephthalic acid.

Still another object of the invention is to provide a bubble columnreactor that facilitates improved liquid-phase oxidation reactions withreduced formation of impurities.

Yet another object of the invention is to provide a more effective andeconomical system for producing pure terephthalic acid (PTA) vialiquid-phase oxidation of para-xylene to produce crude terephthalic acid(CTA) and subsequently, purifying the CTA to PTA.

A further object of the invention is to provide a bubble column reactorfor oxidizing para-xylene and producing a CTA product capable of beingpurified without requiring heat-promoted dissolution of the CTA inwater, hydrogenation of the dissolved CTA, and/or multi-stepcrystallization of the hydrogenated PTA.

It should be noted that the scope of the present invention, as definedin the appended claims, is not limited to processes or apparatusescapable of realizing all of the objects listed above. Rather, the scopeof the claimed invention may encompass a variety of systems that do notaccomplish all or any of the above-listed objects. Additional objectsand advantages of the present invention will be readily apparent to oneskilled in the art upon reviewing the following detailed description andassociated drawings.

SUMMARY OF THE INVENTION

One embodiment of the present invention concerns a process comprisingoxidizing an oxidizable compound in a liquid phase of a reaction mediumcontained in an agitated reactor, wherein the reaction medium has a gashold-up of at least about 0.6 on a time-averaged and volume-averagedbasis.

Another embodiment of the present invention concerns a process forproducing terephthalic acid comprising the following steps: (a)oxidizing para-xylene a liquid phase of a three-phase reaction mediumcontained within a bubble column reactor to thereby form crudeterephthalic acid, wherein the reaction medium has a gas hold-up of atleast about 0.6 on a time-averaged and volume-averaged basis; and (b)oxidizing at least a portion of the crude terephthalic acid in asecondary oxidation reactor to thereby form purer terephthalic acid.

BRIEF DESCRIPTION OF THE DRAWINGS

Preferred embodiments of the invention are described in detail belowwith reference to the attached drawing figures, wherein;

FIG. 1 is a side view of an oxidation reactor constructed in accordancewith one embodiment of the present invention, particularly illustratingthe introduction of feed, oxidant, and reflux streams into the reactor,the presence of a multi-phase reaction medium in the reactor, and thewithdrawal of a gas and a slurry from the top and bottom of the reactor,respectively;

FIG. 2 is an enlarged sectional side view of the bottom of the bubblecolumn reactor taken along line 2-2 in FIG. 3, particularly illustratingthe location and configuration of a oxidant sparger used to introducethe oxidant stream into the reactor;

FIG. 3 is a top view of the oxidant sparger of FIG. 2, particularlyillustrating the oxidant openings in the top of the oxidant sparger;

FIG. 4 is a bottom view of the oxidant sparger of FIG. 2, particularlyillustrating the oxidant openings in the bottom of the oxidant sparger;

FIG. 5 is a sectional side view of the oxidant sparger taken along line5-5 in FIG. 3, particularly illustrating the orientation of the oxidantopenings in the top and bottom of the oxidant sparger;

FIG. 6 is an enlarged side view of the bottom portion of the bubblecolumn reactor, particular illustrating a system for introducing thefeed stream into the reactor at multiple, vertically-space locations;

FIG. 7 is a sectional top view taken along line 7-7 in FIG. 6,particularly illustrating how the feed introduction system shown in FIG.6 distributes the feed stream into in a preferred radial feed zone (FZ)and more than one azimuthal quadrant (Q1, Q2, Q3, Q4);

FIG. 8 is a sectional top view similar to FIG. 7, but illustrating analternative means for discharging the feed stream into the reactor usingbayonet tubes each having a plurality of small feed openings;

FIG. 9 is an isometric view of an alternative system for introducing thefeed stream into the reaction zone at multiple vertically-spacelocations without requiring multiple vessel penetrations, particularlyillustrating that the feed distribution system can be at least partlysupported on the oxidant sparger;

FIG. 10 is a side view of the single-penetration feed distributionsystem and oxidant sparger illustrated in FIG. 9;

FIG. 11 is a sectional top view taken along line 11-11 in FIG. 10 andfurther illustrating the single-penetration feed distribution systemsupported on the oxidant sparger;

FIG. 12 is an isometric view of an alternative oxidant sparger havingall of the oxidant openings located in the bottom of the ring member;

FIG. 13 is a top view of the alternative oxidant sparger of FIG. 12;

FIG. 14 is a bottom view of the alternative oxidant sparger of FIG. 12,particularly illustrating the location of the bottom openings forintroducing the oxidant stream into the reaction zone;

FIG. 15 is a sectional side view of the oxidant sparger taken along line15-15 in FIG. 13, particularly illustrating the orientation of the loweroxidant openings;

FIG. 16 is a side view of a bubble column reactor equipped with aninternal deaeration vessel near the bottom outlet of the reactor;

FIG. 17 is an enlarged sectional side view of the lower portion of thebubble column reactor of FIG. 16 taken along line 17-17 in FIG. 18,particularly illustrating the configuration of the internal deaerationvessel positioned at the bottom outlet of the bubble column reactor;

FIG. 18 is a sectional top view taken along line 18-18 in FIG. 16,particularly illustrating a vortex breaker disposed in the deaerationvessel;

FIG. 19 is a side view of a bubble column reactor equipped with anexternal deaeration vessel and illustrating the manner in which aportion of the deaerated slurry exiting the bottom of the deaerationvessel can be used to flush out a de-inventorying line coupled to thebottom of the reactor;

FIG. 20 is a side view of a bubble column reactor equipped with a hybridinternal/external deaeration vessel for disengaging the gas phase of areaction medium withdrawn from an elevated side location in the reactor;

FIG. 21 is a side view of a bubble column reactor equipped with analternative hybrid deaeration vessel near the bottom of the reactor;

FIG. 22 is an enlarged sectional side view of the lower portion of thebubble column reactor of FIG. 21, particularly illustrating the use ofan alternative oxidant sparger employing inlet conduits that receive theoxidant stream through the bottom head of the reactor;

FIG. 23 is an enlarged sectional side view similar to FIG. 22,particularly illustrating an alternative means for introducing theoxidant stream into the reactor via a plurality of openings in the lowerhead of the reactor and, optionally, employing impingement plates tomore evenly distribute the oxidant stream in the reactor;

FIG. 24 is a side view of a bubble column reactor employing an internalflow conduit to help improve dispersion of an oxidizable compound byrecirculating a portion of the reaction medium from an upper portion ofthe reactor to a lower portion of the reactor;

FIG. 25 is a side view of a bubble column reactor employing an externalflow conduit to help improve dispersion of the oxidizable compound byrecirculating a portion of the reaction medium from an upper portion ofthe reactor to a lower portion of the reactor;

FIG. 26 is a sectional side view of a horizontal eductor that can beused to improve dispersion of the oxidizable compound in an oxidationreactor, particularly illustrating an eductor that uses incoming liquidfeed to draw reaction medium into the eductor and discharges the mixtureof feed and reaction medium into a reaction zone at high velocity;

FIG. 27 is a sectional side view of a vertical eductor that can be usedimprove dispersion of the oxidizable compound in an oxidation reactor,particularly illustrating an eductor that combines the liquid feed andinlet gas and uses the combined two-phase fluid to draw reaction mediuminto the eductor and discharge the mixture of liquid feed, inlet gas,and reaction medium into a reaction zone at high velocity;

FIG. 28 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating the reactionmedium being theoretically partitioned into 30 horizontal slices ofequal volume in order to quantify certain gradients in the reactionmedium;

FIG. 29 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating first and seconddiscrete 20-percent continuous volumes of the reaction medium that havesubstantially different oxygen concentrations and/or oxygen consumptionrates;

FIG. 30 is a side view of two stacked reaction vessels, with or withoutoptional mechanical agitation, containing a multi-phase reaction medium,particularly illustrating that the vessels contain discrete 20-percentcontinuous volumes of the reaction medium having substantially differentoxygen concentrations and/or oxygen consumption rates;

FIG. 31 is a side view of three side-by-side reaction vessels, with orwithout optional mechanical agitation, containing a multi-phase reactionmedium, particularly illustrating that the vessels contain discrete20-percent continuous volumes of the reaction medium havingsubstantially different oxygen concentrations and/or oxygen consumptionrates;

FIG. 32 is a side view of a staged-velocity bubble column reactor havinga broad lower reaction zone and a narrow upper reaction zone;

FIG. 33 is a side view of a bubble column reactor equipped with anupright divider wall for adding upright surface area that contacts thereaction medium;

FIG. 34 is a sectional view taken along line 34-34 in FIG. 33,particularly illustrating that the divider wall is a planar memberdividing the reaction zone into two substantially equal sections;

FIG. 35 is a side view a bubble column reactor equipped with a shortenedupright divider wall for adding upright surface area that contacts thereaction medium;

FIG. 36 is a side view of a bubble column reactor equipped with ashortened and curved upright divider wall for adding upright surfacearea that contacts the reaction medium;

FIG. 37 is a sectional view taken along line 37-37 in FIG. 36,particularly illustrating that the curved upright divider wall is agenerally S-shaped member dividing a portion of the reaction zone intotwo substantially equal sections;

FIG. 38 is a side view of a bubble column reactor equipped with ashortened upright internal member for adding upright surface area thatcontacts the reaction medium;

FIG. 39 is a sectional view taken along line 39-39 in FIG. 38,particularly illustrating that the upright internal member has an “X”shape and the edges of the internal member do not extend all the way tothe reactor side wall;

FIG. 40 is a side view of a bubble column reactor equipped withalternating, differently-configured, upright internal members for addingupright surface area that contacts the reaction medium;

FIG. 41 is a sectional view taken along line 41-41 in FIG. 40,particularly illustrating one configuration of the upright members thathas an “X” shape and divides a portion of the reaction zone into foursubstantially equal quadrants;

FIG. 42 is a sectional view taken along line 42-42 in FIG. 40,particularly illustrating the other configuration of the upright membersthat divides a portion of the reaction zone into eight substantiallyequal wedge-shaped sections;

FIG. 43 is a side view of a bubble column reactor equipped with aplurality of helicoid-shaped internal members for adding upright surfacearea that contacts the reaction medium;

FIG. 44 is a sectional view taken along line 44-44 in FIG. 43,particularly illustrating the shape of one of the helicoid-shapedinternal members;

FIG. 45 is a side view of a bubble column reactor equipped a pluralityof baffles, each comprising a plurality of cylindrical bars forcontacting the reaction medium;

FIG. 46 is an enlarged isometric view of the baffles of FIG. 45,particularly illustrating the manner in which the cylindrical bars ofadjacent baffles are rotated 90 degrees relative to one another;

FIG. 47 is a sectional view taken along line 47-47 in FIG. 45,particularly illustrating a single one of the baffles;

FIG. 48 is a side view of a bubble column reactor equipped with aplurality of baffles, each comprising a plurality of L-section membersfor contacting the reaction medium;

FIG. 49 is an enlarged side view of the baffles of FIG. 48, particularlyillustrating the manner in which the L-section members of adjacentbaffles are rotated 90 degrees relative to one another;

FIG. 50 is a sectional view taken along line 50-50 in FIG. 48,particularly illustrating a single one of the baffles;

FIG. 51 is a side view of a bubble column reactor equipped with a singlemonolithic, cylindrical, diamond-shaped baffle for contacting thereaction medium;

FIG. 52 is an enlarged side view of the monolithic baffle of FIG. 51;

FIG. 53 is a sectional view taken along line 53-53 in FIG. 51 andillustrating the cylindrical nature of the monolithic baffle;

FIGS. 54A and 54B are magnified views of crude terephthalic acid (CTA)particles produced in accordance with one embodiment of the presentinvention, particularly illustrating that each CTA particle is a lowdensity, high surface area particle composed of a plurality ofloosely-bound CTA sub-particles;

FIG. 55A and 55B are magnified views of a conventionally-produced CTA,particularly illustrating that the conventional CTA particle has alarger particle size, lower density, and lower surface area than theinventive CTA particle of FIGS. 54A and 54B;

FIG. 56 is a simplified process flow diagram of a prior art process formaking purified terephthalic acid (PTA);

FIG. 57 is a simplified process flow diagram of a process for making PTAin accordance with one embodiment of the present invention; and

FIG. 58 is a table summarizing various operating parameters of a bubblecolumn oxidation reactor, wherein certain of the operating parametershave been adjusted in accordance with the description provide in theExamples section.

DETAILED DESCRIPTION

One embodiment of the present invention concerns the liquid-phasepartial oxidation of an oxidizable compound. Such oxidation ispreferably carried out in the liquid phase of a multi-phase reactionmedium contained in one or more agitated reactors. Suitable agitatedreactors include, for example, bubble-agitated reactors (e.g., bubblecolumn reactors), mechanically agitated reactors (e.g., continuousstirred tank reactors), and flow agitated reactors (e.g., jet reactors).In one embodiment of the invention, the liquid-phase oxidation iscarried out in a single bubble column reactor.

As used herein, the term “bubble column reactor” shall denote a reactorfor facilitating chemical reactions in a multi-phase reaction medium,wherein agitation of the reaction medium is provided primarily by theupward movement of gas bubbles through the reaction medium. As usedherein, the term “agitation” shall denote work dissipated into thereaction medium causing fluid flow and/or mixing. As used herein, theterms “majority”, “primarily”, and “predominately” shall mean more than50 percent. As used herein, the term “mechanical agitation” shall denoteagitation of the reaction medium caused by physical movement of a rigidor flexible element(s) against or within the reaction medium. Forexample, mechanical agitation can be provided by rotation, oscillation,and/or vibration of internal stirrers, paddles, vibrators, or acousticaldiaphragms located in the reaction medium. As used herein, the term“flow agitation” shall denote agitation of the reaction medium caused byhigh velocity injection and/or recirculation of one or more fluids inthe reaction medium. For example, flow agitation can be provided bynozzles, ejectors, and/or eductors.

In a preferred embodiment of the present invention, less than about 40percent of the agitation of the reaction medium in the bubble columnreactor during oxidation is provided by mechanical and/or flowagitation, more preferably less than about 20 percent of the agitationis provided by mechanical and/or flow agitation, and most preferablyless than 5 percent of the agitation is provided by mechanical and/orflow agitation. Preferably, the amount of mechanical and/or flowagitation imparted to the multi-phase reaction medium during oxidationis less than about 3 kilowatts per cubic meter of the reaction medium,more preferably less than about 2 kilowatts per cubic meter, and mostpreferably less than 1 kilowatt per cubic meter.

Referring now to FIG. 1, a preferred bubble column reactor 20 isillustrated as comprising a vessel shell 22 having of a reaction section24 and a disengagement section 26. Reaction section 24 defines aninternal reaction zone 28, while disengagement section 26 defines aninternal disengagement zone 30. A predominately liquid-phase feed streamis introduced into reaction zone 28 via feed inlets 32 a,b,c,d. Apredominately gas-phase oxidant stream is introduced into reaction zone28 via an oxidant sparger 34 located in the lower portion of reactionzone 28. The liquid-phase feed stream and gas-phase oxidant streamcooperatively form a multi-phase reaction medium 36 within reaction zone28. Multi-phase reaction medium 36 comprises a liquid phase and a gasphase. More preferably, multiphase reaction medium 36 comprises athree-phase medium having solid-phase, liquid-phase, and gas-phasecomponents. The solid-phase component of the reaction medium 36preferably precipitates within reaction zone 28 as a result of theoxidation reaction carried out in the liquid phase of reaction medium36. Bubble column reactor 20 includes a slurry outlet 38 located nearthe bottom of reaction zone 28 and a gas outlet 40 located near the topof disengagement zone 30. A slurry effluent comprising liquid-phase andsolid-phase components of reaction medium 36 is withdrawn from reactionzone 28 via slurry outlet 38, while a predominantly gaseous effluent iswithdrawn from disengagement zone 30 via gas outlet 40.

The liquid-phase feed stream introduced into bubble column reactor 20via feed inlets 32 a,b,c,d preferably comprises an oxidizable compound,a solvent, and a catalyst system.

The oxidizable compound present in the liquid-phase feed streampreferably comprises at least one hydrocarbyl group. More preferably,the oxidizable compound is an aromatic compound. Still more preferably,the oxidizable compound is an aromatic compound with at least oneattached hydrocarbyl group or at least one attached substitutedhydrocarbyl group or at least one attached heteroatom or at least oneattached carboxylic acid function (—COOH). Even more preferably, theoxidizable compound is an aromatic compound with at least one attachedhydrocarbyl group or at least one attached substituted hydrocarbyl groupwith each attached group comprising from 1 to 5 carbon atoms. Yet stillmore preferably, the oxidizable compound is an aromatic compound havingexactly two attached groups with each attached group comprising exactlyone carbon atom and consisting of methyl groups and/or substitutedmethyl groups and/or at most one carboxylic acid group. Even still morepreferably, the oxidizable compound is para-xylene, meta-xylene,para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluicacid, and/or acetaldehyde. Most preferably, the oxidizable compound ispara-xylene.

A “hydrocarbyl group”, as defined herein, is at least one carbon atomthat is bonded only to hydrogen atoms or to other carbon atoms. A“substituted hydrocarbyl group”, as defined herein, is at least onecarbon atom bonded to at least one heteroatom and to at least onehydrogen atom. “Heteroatoms”, as defined herein, are all atoms otherthan carbon and hydrogen atoms. Aromatic compounds, as defined herein,comprise an aromatic ring, preferably having at least 6 carbon atoms,even more preferably having only carbon atoms as part of the ring.Suitable examples of such aromatic rings include, but are not limitedto, benzene, biphenyl, terphenyl, naphthalene, and other carbon-basedfused aromatic rings.

Suitable examples of the oxidizable compound include aliphatichydrocarbons (e.g., alkanes, branched alkanes, cyclic alkanes, aliphaticalkenes, branched alkenes, and cyclic alkenes); aliphatic aldehydes(e.g., acetaldehyde, propionaldehyde, isobutyraldehyde, andn-butyraldehyde); aliphatic alcohols (e.g., ethanol, isopropanol,n-propanol, n-butanol, and isobutanol); aliphatic ketones (e.g.,dimethyl ketone, ethyl methyl ketone, diethyl ketone, and isopropylmethyl ketone); aliphatic esters (e.g., methyl formate, methyl acetate,ethyl acetate); aliphatic peroxides, peracids, and hydroperoxides (e.g.,t-butyl hydroperoxide, peracetic acid, and di-t-butyl hydroperoxide);aliphatic compounds with groups that are combinations of the abovealiphatic species plus other heteroatoms (e.g., aliphatic compoundscomprising one or more molecular segments of hydrocarbons, aldehydes,alcohols, ketones, esters, peroxides, peracids, and/or hydroperoxides incombination with sodium, bromine, cobalt, manganese, and zirconium);various benzene rings, naphthalene rings, biphenyls, terphenyls, andother aromatic groups with one or more attached hydrocarbyl groups(e.g., toluene, ethylbenzene, isopropylbenzene, n-propylbenzene,neopentylbenzene, para-xylene, meta-xylene, ortho-xylene, all isomers oftrimethylbenzenes, all isomers of tetramethylbenzenes,pentamethylbenzene, hexamethylbenzene, all isomers ofethyl-methylbenzenes, all isomers of diethylbenzenes, all isomers ofethyl-dimethylbenzenes, all isomers of dimethylnaphthalenes, all isomersof ethyl-methylnaphthalenes, all isomers of diethylnaphthalenes allisomers of dimethylbiphenyls, all isomers of ethyl-methylbiphenyls, andall isomers of diethylbiphenyls, stilbene and with one or more attachedhydrocarbyl groups, fluorene and with one or more attached hydrocarbylgroups, anthracene and with one or more attached hydrocarbyl groups, anddiphenylethane and with one or more attached hydrocarbyl groups);various benzene rings, naphthalene rings, biphenyls, terphenyls, andother aromatic groups with one or more attached hydrocarbyl groupsand/or one or more attached heteroatoms, which may connect to otheratoms or groups of atoms (e.g., phenol, all isomers of methylphenols,all isomers of dimethylphenols, all isomers of naphthols, benzyl methylether, all isomers of bromophenols, bromobenzene, all isomers ofbromotoluenes including alpha-bromotoluene, dibromobenzene, cobaltnaphthenate, and all isomers of bromobiphenyls); various benzene rings,naphthalene rings, biphenyls, terphenyls, and other aromatic groups withone or more attached hydrocarbyl groups and/or one or more attachedheteroatoms and/or one or more attached substituted hydrocarbyl groups(e.g., benzaldehyde, all isomers of bromobenzaldehydes, all isomers ofbrominated tolualdehydes including all isomers ofalpha-bromotolualdehydes, all isomers of hydroxybenzaldehydes, allisomers of bromo-hydroxybenzaldehydes, all isomers of benzenedicarboxaldehydes, all isomers of benzene tricarboxaldehydes,para-tolualdehyde, meta-tolualdehyde, ortho-tolualdehyde, all isomers oftoluene dicarboxaldehydes, all isomers of toluene tricarboxaldehydes,all isomers of toluene tetracarboxaldehydes, all isomers ofdimethylbenzene dicarboxaldehydes, all isomers of dimethylbenzenetricarboxaldehydes, all isomers of dimethylbenzene tetracarboxaldehydes,all isomers of trimethylbenzene tricarboxaldehydes, all isomers ofethyltolualdehydes, all isomers of trimethylbenzene dicarboxaldehydes,tetramethylbenzene dicarboxaldehyde, hydroxymethyl-benzene, all isomersof hydroxymethyl-toluenes, all isomers of hydroxymethyl-bromotoluenes,all isomers of hydroxymethyl-tolualdehydes, all isomers ofhydroxymethyl-bromotolualdehydes, benzyl hydroperoxide, benzoylhydroperoxide, all isomers of tolyl methyl-hydroperoxides, and allisomers of methylphenol methyl-hydroperoxides); various benzene rings,naphthalenes rings, biphenyls, terphenyls, and other aromatic groupswith one or more attached selected groups, selected groups meaninghydrocarbyl groups and/or attached heteroatoms and/or substitutedhydrocarbyl groups and/or carboxylic acid groups and/or peroxy acidgroups (e.g., benzoic acid, para-toluic acid, meta-toluic acid,ortho-toluic acid, all isomers of ethylbenzoic acids, all isomers ofpropylbenzoic acids, all isomers of butylbenzoic acids, all isomers ofpentylbenzoic acids, all isomers of dimethylbenzoic acids, all isomersof ethylmethylbenzoic acids, all isomers of trimethylbenzoic acids, allisomers of tetramethylbenzoic acids, pentamethylbenzoic acid, allisomers of diethylbenzoic acids, all isomers of benzene dicarboxylicacids, all isomers of benzene tricarboxylic acids, all isomers ofmethylbenzene dicarboxylic acids, all isomers of dimethylbenzenedicarboxylic acids, all isomers of methylbenzene tricarboxylic acids,all isomers of bromobenzoic acids, all isomers of dibromobenzoic acids,all isomers of bromotoluic acids including alpha-bromotoluic acids,tolyl acetic acid, all isomers of hydroxybenzoic acids, all isomers ofhydroxymethyl-benzoic acids, all isomers of hydroxytoluic acids, allisomers of hydroxymethyl-toluic acids, all isomers ofhydroxymethyl-benzene dicarboxylic acids, all isomers ofhydroxybromobenzoic acids, all isomers of hydroxybromotoluic acids, allisomers of hydroxymethyl-bromobenzoic acids, all isomers of carboxybenzaldehydes, all isomers of dicarboxy benzaldehydes, perbenzoic acid,all isomers of hydroperoxymethyl-benzoic acids, all isomers ofhydroperoxymethyl-hydroxybenzoic acids, all isomers ofhydroperoxycarbonyl-benzoic acids, all isomers ofhydroperoxycarbonyl-toluenes, all isomers of methylbiphenyl carboxylicacids, all isomers of dimethylbiphenyl carboxylic acids, all isomers ofmethylbiphenyl dicarboxylic acids, all isomers of biphenyl tricarboxylicacids, all isomers of stilbene with one or more attached selectedgroups, all isomers of fluorenone with one or more attached selectedgroups, all isomers of naphthalene with one or more attached selectedgroups, benzil, all isomers of benzil with one or more attached selectedgroups, benzophenone, all isomers of benzophenone with one or moreattached selected groups, anthraquinone, all isomers of anthraquinonewith one or more attached selected groups, all isomers of diphenylethanewith one or more attached selected groups, benzocoumarin, and allisomers of benzocoumarin with one or more attached selected groups).

If the oxidizable compound present in the liquid-phase feed stream is anormally-solid compound (i.e., is a solid at standard temperature andpressure), it is preferred for the oxidizable compound to besubstantially dissolved in the solvent when introduced into reactionzone 28. It is preferred for the boiling point of the oxidizablecompound at atmospheric pressure to be at least about 50° C. Morepreferably, the boiling point of the oxidizable compound is in the rangeof from about 80 to about 400° C., and most preferably in the range offrom 125 to 155° C. The amount of oxidizable compound present in theliquid-phase feed is preferably in the range of from about 2 to about 40weight percent, more preferably in the range of from about 4 to about 20weight percent, and most preferably in the range of from 6 to 15 weightpercent.

It is now noted that the oxidizable compound present in the liquid-phasefeed may comprise a combination of two or more different oxidizablechemicals. These two or more different chemical materials can be fedcommingled in the liquid-phase feed stream or may be fed separately inmultiple feed streams. For example, an oxidizable compound comprisingpara-xylene, meta-xylene, para-tolualdehyde, para-toluic acid, andacetaldehyde may be fed to the reactor via a single inlet or multipleseparate inlets.

The solvent present in the liquid-phase feed stream preferably comprisesan acid component and a water component. The solvent is preferablypresent in the liquid-phase feed stream at a concentration in the rangeof from about 60 to about 98 weight percent, more preferably in therange of from about 80 to about 96 weight percent, and most preferablyin the range of from 85 to 94 weight percent. The acid component of thesolvent is preferably primarily an organic low molecular weightmonocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbonatoms. Most preferably, the acid component of the solvent is primarilyacetic acid. Preferably, the acid component makes up at least about 75weight percent of the solvent, more preferably at least about 80 weightpercent of the solvent, and most preferably 85 to 98 weight percent ofthe solvent, with the balance being primarily water. The solventintroduced into bubble column reactor 20 can include small quantities ofimpurities such as, for example, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha-bromo-para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. It is preferred that the total amount of impurities in thesolvent introduced into bubble column reactor 20 is less than about 3weight percent.

The catalyst system present in the liquid-phase feed stream ispreferably a homogeneous, liquid-phase catalyst system capable ofpromoting oxidation (including partial oxidation) of the oxidizablecompound. More preferably, the catalyst system comprises at least onemultivalent transition metal. Still more preferably, the multivalenttransition metal comprises cobalt. Even more preferably, the catalystsystem comprises cobalt and bromine. Most preferably, the catalystsystem comprises cobalt, bromine, and manganese.

When cobalt is present in the catalyst system, it is preferred for theamount of cobalt present in the liquid-phase feed stream to be such thatthe concentration of cobalt in the liquid phase of reaction medium 36 ismaintained in the range of from about 300 to about 6,000 parts permillion by weight (ppmw), more preferably in the range of from about 700to about 4,200 ppmw, and most preferably in the range of from 1,200 to3,000 ppmw. When bromine is present in the catalyst system, it ispreferred for the amount of bromine present in the liquid-phase feedstream to be such that the concentration of bromine in the liquid phaseof reaction medium 36 is maintained in the range of from about 300 toabout 5,000 ppmw, more preferably in the range of from about 600 toabout 4,000 ppmw, and most preferably in the range of from 900 to 3,000ppmw. When manganese is present in the catalyst system, it is preferredfor the amount of manganese present in the liquid-phase feed stream tobe such that the concentration of manganese in the liquid phase ofreaction medium 36 is maintained in the range of from about 20 to about1000 ppmw, more preferably in the range of from about 40 to about 500ppmw, most preferably in the range of from 50 to 200 ppmw.

The concentrations of the cobalt, bromine, and/or manganese in theliquid phase of reaction medium 36, provided above, are expressed on atime-averaged and volume-averaged basis. As used herein, the term“time-averaged” shall denote an average of at least 10 measurementstaken equally over a continuous period of at least 100 seconds. As usedherein, the term “volume-averaged” shall denote an average of at least10 measurements taken at uniform 3-dimensional spacing throughout acertain volume.

The weight ratio of cobalt to bromine (Co:Br) in the catalyst systemintroduced into reaction zone 28 is preferably in the range of fromabout 0.25:1 to about 4:1, more preferably in the range of from about0.5:1 to about 3:1, and most preferably in the range of from 0.75:1 to2:1. The weight ratio of cobalt to manganese (Co:Mn) in the catalystsystem introduced into reaction zone 28 is preferably in the range offrom about 0.3:1 to about 40:1, more preferably in the range of fromabout 5:1 to about 30:1, and most preferably in the range of from 10:1to 25:1.

The liquid-phase feed stream introduced into bubble column reactor 20can include small quantities of impurities such as, for example,toluene, ethylbenzene, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha bromo para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. When bubble column reactor 20 is employed for theproduction of terephthalic acid, meta-xylene and ortho-xylene are alsoconsidered impurities. It is preferred that the total amount ofimpurities in the liquid-phase feed stream introduced into bubble columnreactor 20 is less than about 3 weight percent.

Although FIG. 1 illustrates an embodiment where the oxidizable compound,the solvent, and the catalyst system are mixed together and introducedinto bubble column reactor 20 as a single feed stream, in an alternativeembodiment of the present invention, the oxidizable compound, thesolvent, and the catalyst can be separately introduced into bubblecolumn reactor 20. For example, it is possible to feed a purepara-xylene stream into bubble column reactor 20 via an inlet separatefrom the solvent and catalyst inlet(s).

The predominately gas-phase oxidant stream introduced into bubble columnreactor 20 via oxidant sparger 34 comprises molecular oxygen (O₂).Preferably, the oxidant stream comprises in the range of from about 5 toabout 40 mole percent molecular oxygen, more preferably in the range offrom about 15 to about 30 mole percent molecular oxygen, and mostpreferably in the range of from 18 to 24 mole percent molecular oxygen.It is preferred for the balance of the oxidant stream to be comprisedprimarily of a gas or gasses, such as nitrogen, that are inert tooxidation. More preferably, the oxidant stream consists essentially ofmolecular oxygen and nitrogen. Most preferably, the oxidant stream isdry air that comprises about 21 mole percent molecular oxygen and about78 to about 81 mole percent nitrogen. In an alternative embodiment ofthe present invention, the oxidant stream can comprise substantiallypure oxygen.

Referring again to FIG. 1, bubble column reactor 20 is preferablyequipped with a reflux distributor 42 positioned above an upper surface44 of reaction medium 36. Reflux distributor 42 is operable to introducedroplets of a predominately liquid-phase reflux stream intodisengagement zone 30 by any means of droplet formation known in theart. More preferably, reflux distributor 42 produces a spray of dropletsdirected downwardly towards upper surface 44 of reaction medium 36.Preferably, this downward spray of droplets affects (i.e., engages andinfluences) at least about 50 percent of the maximum horizontalcross-sectional area of disengagement zone 30. More preferably, thespray of droplets affects at least about 75 percent of the maximumhorizontal cross-sectional area of disengagement zone 30. Mostpreferably, the spray of droplets affects at least 90 percent of themaximum horizontal cross-sectional area of disengagement zone 30. Thisdownward liquid reflux spray can help prevent foaming at or above uppersurface 44 of reaction medium 36 and can also aid in the disengagementof any liquid or slurry droplets entrained in the upwardly moving gasthat flows towards gas outlet 40. Further, the liquid reflux may serveto reduce the amount of particulates and potentially precipitatingcompounds (e.g., dissolved benzoic acid, para-toluic acid, 4-CBA,terephthalic acid, and catalyst metal salts) exiting in the gaseouseffluent withdrawn from disengagement zone 30 via gas outlet 40. Inaddition, the introduction of reflux droplets into disengagement zone 30can, by a distillation action, be used to adjust the composition of thegaseous effluent withdrawn via gas outlet 40.

The liquid reflux stream introduced into bubble column reactor 20 viareflux distributor 42 preferably has about the same composition as thesolvent component of the liquid-phase feed stream introduced into bubblecolumn reactor 20 via feed inlets 32 a,b,c,d. Thus, it is preferred forthe liquid reflux stream to comprise an acid component and water. Theacid component of the reflux stream is preferably a low molecular weightorganic monocarboxylic acid having 1-6 carbon atoms, more preferably 2carbon atoms. Most preferably, the acid component of the reflux streamis acetic acid. Preferably, the acid component makes up at least about75 weight percent of the reflux stream, more preferably at least about80 weight percent of the reflux stream, and most preferably 85 to 98weight percent of the reflux stream, with the balance being water.Because the reflux stream typically has substantially the samecomposition as the solvent in the liquid-phase feed stream, when thisdescription refers to the “total solvent” introduced into the reactor,such “total solvent” shall include both the reflux stream and thesolvent portion of the feed stream.

During liquid-phase oxidation in bubble column reactor 20, it ispreferred for the feed, oxidant, and reflux streams to be substantiallycontinuously introduced into reaction zone 28, while the gas and slurryeffluent streams are substantially continuously withdrawn from reactionzone 28. As used herein, the term “substantially continuously” shallmean for a period of at least 10 hours interrupted by less than 10minutes. During oxidation, it is preferred for the oxidizable compound(e.g., para-xylene) to be substantially continuously introduced intoreaction zone 28 at a rate of at least about 8,000 kilograms per hour,more preferably at a rate in the range of from about 13,000 to about80,000 kilograms per hour, still more preferably in the range of fromabout 18,000 to about 50,000 kilograms per hour, and most preferably inthe range of from 22,000 to 30,000 kilograms per hour. Although it isgenerally preferred for the flow rates of the incoming feed, oxidant,and reflux streams to be substantially steady, it is now noted that oneembodiment of the presenting invention contemplates pulsing the incomingfeed, oxidant, and/or reflux stream in order to improve mixing and masstransfer. When the incoming feed, oxidant, and/or reflux stream areintroduced in a pulsed fashion, it is preferred for their flow rates tovary within about 0 to about 500 percent of the steady-state flow ratesrecited herein, more preferably within about 30 to about 200 percent ofthe steady-state flow rates recited herein, and most preferably within80 to 120 percent of the steady-state flow rates recited herein.

The average space-time rate of reaction (STR) in bubble column oxidationreactor 20 is defined as the mass of the oxidizable compound fed perunit volume of reaction medium 36 per unit time (e.g., kilograms ofpara-xylene fed per cubic meter per hour). In conventional usage, theamount of oxidizable compound not converted to product would typicallybe subtracted from the amount of oxidizable compound in the feed streambefore calculating the STR. However, conversions and yields aretypically high for many of the oxidizable compounds preferred herein(e.g., para-xylene), and it is convenient to define the term herein asstated above. For reasons of capital cost and operating inventory, amongothers, it is generally preferred that the reaction be conducted with ahigh STR. However, conducting the reaction at increasingly higher STRmay affect the quality or yield of the partial oxidation. Bubble columnreactor 20 is particularly useful when the STR of the oxidizablecompound (e.g., para-xylene) is in the range of from about 25 kilogramsper cubic meter per hour to about 400 kilograms per cubic meter perhour, more preferably in the range of from about 30 kilograms per cubicmeter per hour to about 250 kilograms per cubic meter per hour, stillmore preferably from about 35 kilograms per cubic meter per hour toabout 150 kilograms per cubic meter per hour, and most preferably in therange of from 40 kilograms per cubic meter per hour to 100 kilograms percubic meter per hour.

The oxygen-STR in bubble column oxidation reactor 20 is defined as theweight of molecular oxygen consumed per unit volume of reaction medium36 per unit time (e.g., kilograms of molecular oxygen consumed per cubicmeter per hour). For reasons of capital cost and oxidative consumptionof solvent, among others, it is generally preferred that the reaction beconducted with a high oxygen-STR. However, conducting the reaction atincreasingly higher oxygen-STR eventually reduces the quality or yieldof the partial oxidation. Without being bound by theory, it appears thatthis possibly relates to the transfer rate of molecular oxygen from thegas phase into the liquid at the interfacial surface area and thenceinto the bulk liquid. Too high an oxygen-STR possibly leads to too low adissolved oxygen content in the bulk liquid phase of the reactionmedium.

The global-average-oxygen-STR is defined herein as the weight of alloxygen consumed in the entire volume of reaction medium 36 per unit time(e.g., kilograms of molecular oxygen consumed per cubic meter per hour).Bubble column reactor 20 is particularly useful when theglobal-average-oxygen-STR is in the range of from about 25 kilograms percubic meter per hour to about 400 kilograms per cubic meter per hour,more preferably in the range of from about 30 kilograms per cubic meterper hour to about 250 kilograms per cubic meter per hour, still morepreferably from about 35 kilograms per cubic meter per hour to about 150kilograms per cubic meter per hour, and most preferably in the range offrom 40 kilograms per cubic meter per hour to 100 kilograms per cubicmeter per hour.

During oxidation in bubble column reactor 20, it is preferred for theratio of the mass flow rate of the total solvent (from both the feed andreflux streams) to the mass flow rate of the oxidizable compoundentering reaction zone 28 to be maintained in the range of from about2:1 to about 50:1, more preferably in the range of from about 5:1 toabout 40:1, and most preferably in the range of from 7.5:1 to 25:1.Preferably, the ratio of the mass flow rate of solvent introduced aspart of the feed stream to the mass flow rate of solvent introduced aspart of the reflux stream is maintained in the range of from about 0.5:1to no reflux stream flow whatsoever, more preferably in the range offrom about 0.5:1 to about 4:1, still more preferably in the range offrom about 1:1 to about 2:1, and most preferably in the range of from1.25:1 to 1.5:1.

During liquid-phase oxidation in bubble column reactor 20, it ispreferred for the oxidant stream to be introduced into bubble columnreactor 20 in an amount that provides molecular oxygen somewhatexceeding the stoichiometric oxygen demand. The amount of excessmolecular oxygen required for best results with a particular oxidizablecompound affects the overall economics of the liquid-phase oxidation.During liquid-phase oxidation in bubble column reactor 20, it ispreferred that the ratio of the mass flow rate of the oxidant stream tothe mass flow rate of the oxidizable organic compound (e.g.,para-xylene) entering reactor 20 is maintained in the range of fromabout 0.5:1 to about 20:1, more preferably in the range of from about1:1 to about 10:1, and most preferably in the range of from 2:1 to 6:1.

Referring again to FIG. 1, the feed, oxidant, and reflux streamsintroduced into bubble column reactor 20 cooperatively form at least aportion of multi-phase reaction medium 36. Reaction medium 36 ispreferably a three-phase medium comprising a solid phase, a liquidphase, and a gas phase. As mentioned above, oxidation of the oxidizablecompound (e.g., para-xylene) takes place predominately in the liquidphase of reaction medium 36. Thus, the liquid phase of reaction medium36 comprises dissolved oxygen and the oxidizable compound. Theexothermic nature of the oxidation reaction that takes place in bubblecolumn reactor 20 causes a portion of the solvent (e.g., acetic acid andwater) introduced via feed inlets 32 a,b,c,d to boil/vaporize. Thus, thegas phase of reaction medium 36 in reactor 20 is formed primarily ofvaporized solvent and an undissolved, unreacted portion of the oxidantstream. Certain prior art oxidation reactors employ heat exchangetubes/fins to heat or cool the reaction medium. However, such heatexchange structures may be undesirable in the inventive reactor andprocess described herein. Thus, it is preferred for bubble columnreactor 20 to include substantially no surfaces that contact reactionmedium 36 and exhibit a time-averaged heat flux greater than 30,000watts per meter squared.

The concentration of dissolved oxygen in the liquid phase of reactionmedium 36 is a dynamic balance between the rate of mass transfer fromthe gas phase and the rate of reactive consumption within the liquidphase (i.e. it is not set simply by the partial pressure of molecularoxygen in the supplying gas phase, though this is one factor in thesupply rate of dissolved oxygen and it does affect the limiting upperconcentration of dissolved oxygen). The amount of dissolved oxygenvaries locally, being higher near bubble interfaces. Globally, theamount of dissolved oxygen depends on the balance of supply and demandfactors in different regions of reaction medium 36. Temporally, theamount of dissolved oxygen depends on the uniformity of gas and liquidmixing relative to chemical consumption rates. In designing to matchappropriately the supply of and demand for dissolved oxygen in theliquid phase of reaction medium 36, it is preferred for thetime-averaged and volume-averaged oxygen concentration in the liquidphase of reaction medium 36 to be maintained above about 1 ppm molar,more preferably in the range from about 4 to about 1,000 ppm molar,still more preferably in the range from about 8 to about 500 ppm molar,and most preferably in the range from 12 to 120 ppm molar.

The liquid-phase oxidation reaction carried out in bubble column reactor20 is preferably a precipitating reaction that generates solids. Morepreferably, the liquid-phase oxidation carried out in bubble columnreactor 20 causes at least about 10 weight percent of the oxidizablecompound (e.g., para-xylene) introduced into reaction zone 28 to form asolid compound (e.g., crude terephthalic acid particles) in reactionmedium 36. Still more preferably, the liquid-phase oxidation causes atleast about 50 weight percent of the oxidizable compound to form a solidcompound in reaction medium 36. Most preferably, the liquid-phaseoxidation causes at least 90 weight percent of the oxidizable compoundto form a solid compound in reaction medium 36. It is preferred for thetotal amount of solids in reaction medium 36 to be greater than about 3percent by weight on a time-averaged and volume-averaged basis. Morepreferably, the total amount of solids in reaction medium 36 ismaintained in the range of from about 5 to about 40 weight percent,still more preferably in the range of from about 10 to about 35 weightpercent, and most preferably in the range of from 15 to 30 weightpercent. It is preferred for a substantial portion of the oxidationproduct (e.g., terephthalic acid) produced in bubble column reactor 20to be present in reaction medium 36 as solids, as opposed to remainingdissolved in the liquid phase of reaction medium 36. The amount of thesolid phase oxidation product present in reaction medium 36 ispreferably at least about 25 percent by weight of the total oxidationproduct (solid and liquid phase) in reaction medium 36, more preferablyat least about 75 percent by weight of the total oxidation product inreaction medium 36, and most preferably at least 95 percent by weight ofthe total oxidation product in reaction medium 36. The numerical rangesprovided above for the amount of solids in reaction medium 36 apply tosubstantially steady-state operation of bubble column 20 over asubstantially continuous period of time, not to start-up, shut-down, orsub-optimal operation of bubble column reactor 20. The amount of solidsin reaction medium 36 is determined by a gravimetric method. In thisgravimetric method, a representative portion of slurry is withdrawn fromthe reaction medium and weighed. At conditions that effectively maintainthe overall solid-liquid partitioning present within the reactionmedium, free liquid is removed from the solids portion by sedimentationor filtration, effectively without loss of precipitated solids and withless than about 10 percent of the initial liquid mass remaining with theportion of solids. The remaining liquid on the solids is evaporated todryness, effectively without sublimation of solids. The remainingportion of solids is weighed. The ratio of the weight of the portion ofsolids to the weight of the original portion of slurry is the fractionof solids, typically expressed as a percentage.

The precipitating reaction carried out in bubble column reactor 20 cancause fouling (i.e., solids build-up) on the surface of certain rigidstructures that contact reaction medium 36. Thus, in one embodiment ofthe present invention, it is preferred for bubble column reactor 20 toinclude substantially no internal heat exchange, stirring, or bafflingstructures in reaction zone 28 because such structures would be prone tofouling. If internal structures are present in reaction zone 28, it isdesirable to avoid internal structures having outer surfaces thatinclude a significant amount of upwardly facing planar surface areabecause such upwardly facing planar surfaces would be highly prone tofouling. Thus, if any internal structures are present in reaction zone28, it is preferred for less than about 20 percent of the total upwardlyfacing exposed outer surface area of such internal structures to beformed by substantially planar surfaces inclined less than about 15degrees from horizontal.

Referring again to FIG. 1, the physical configuration of bubble columnreactor 20 helps provide for optimized oxidation of the oxidizablecompound (e.g., para-xylene) with minimal impurity generation. It ispreferred for elongated reaction section 24 of vessel shell 22 toinclude a substantially cylindrical main body 46 and a lower head 48.The upper end of reaction zone 28 is defined by a horizontal plane 50extending across the top of cylindrical main body 46. A lower end 52 ofreaction zone 28 is defined by the lowest internal surface of lower head48. Typically, lower end 52 of reaction zone 28 is located proximate theopening for slurry outlet 38. Thus, elongated reaction zone 28 definedwithin bubble column reactor 20 has a maximum length “L” measured fromthe top end 50 to the bottom end 52 of reaction zone 28 along the axisof elongation of cylindrical main body 46. The length “L” of reactionzone 28 is preferably in the range of from about 10 to about 100 meters,more preferably in the range of from about 20 to about 75 meters, andmost preferably in the range of from 25 to 50 meters. Reaction zone 28has a maximum diameter (width) “D” that is typically equal to themaximum internal diameter of cylindrical main body 46. The maximumdiameter “D” of reaction zone 28 is preferably in the range of fromabout 1 to about 12 meters, more preferably in the range of from about 2to about 10 meters, still more preferably in the range of from about 3.1to about 9 meters, and most preferably in the range of from 4 to 8meters. In a preferred embodiment of the present invention, reactionzone 28 has a length-to-diameter “L:D” ratio in the range of from about6:1 to about 30:1. Still more preferably, reaction zone 28 has an L:Dratio in the range of from about 8:1 to about 20:1. Most preferably,reaction zone 28 has an L:D ratio in the range of from 9:1 to 15:1.

As discussed above, reaction zone 28 of bubble column reactor 20receives multi-phase reaction medium 36. Reaction medium 36 has a bottomend coincident with lower end 52 of reaction zone 28 and a top endlocated at upper surface 44. Upper surface 44 of reaction medium 36 isdefined along a horizontal plane that cuts through reaction zone 28 at avertical location where the contents of reaction zone 28 transitionsfrom a gas-phase-continuous state to a liquid-phase-continuous state.Upper surface 44 is preferably positioned at the vertical location wherethe local time-averaged gas hold-up of a thin horizontal slice of thecontents of reaction zone 28 is 0.9.

Reaction medium 36 has a maximum height “H” measured between its upperand lower ends. The maximum width “W” of reaction medium 36 is typicallyequal to the maximum diameter “D” of cylindrical main body 46. Duringliquid-phase oxidation in bubble column reactor 20, it is preferred thatH is maintained at about 60 to about 120 percent of L, more preferablyabout 80 to about 110 percent of L, and most preferably 85 to 100percent of L. In a preferred embodiment of the present invention,reaction medium 36 has a height-to-width “H:W” ratio greater than about3:1. More preferably, reaction medium 36 has an H:W ratio in the rangeof from about 7:1 to about 25:1. Still more preferably, reaction medium36 has an H:W ratio in the range of from about 8:1 to about 20:1. Mostpreferably, reaction medium 36 has an H:W ratio in the range of from 9:1to 15:1. In one embodiment of the invention, L=H and D=W so that variousdimensions or ratios provide herein for L and D also apply to H and W,and vice-versa.

The relatively high L:D and H:W ratios provided in accordance with anembodiment of the invention can contribute to several importantadvantages of the inventive system. As discussed in further detailbelow, it has been discovered that higher L:D and H:W ratios, as well ascertain other features discussed below, can promote beneficial verticalgradients in the concentrations of molecular oxygen and/or theoxidizable compound (e.g., para-xylene) in reaction medium 36. Contraryto conventional wisdom, which would favor a well-mixed reaction mediumwith relatively uniform concentrations throughout, it has beendiscovered that the vertical staging of the oxygen and/or the oxidizablecompound concentrations facilitates a more effective and economicaloxidation reaction. Minimizing the oxygen and oxidizable compoundconcentrations near the top of reaction medium 36 can help avoid loss ofunreacted oxygen and unreacted oxidizable compound through upper gasoutlet 40. However, if the concentrations of oxidizable compound andunreacted oxygen are low throughout reaction medium 36, then the rateand/or selectivity of oxidation are reduced. Thus, it is preferred forthe concentrations of molecular oxygen and/or the oxidizable compound tobe significantly higher near the bottom of reaction medium 36 than nearthe top of reaction medium 36.

In addition, high L:D and H:W ratios cause the pressure at the bottom ofreaction medium 36 to be substantially greater than the pressure at thetop of reaction medium 36. This vertical pressure gradient is a resultof the height and density of reaction medium 36. One advantage of thisvertical pressure gradient is that the elevated pressure at the bottomof the vessel drives more oxygen solubility and mass transfer than wouldotherwise be achievable at comparable temperatures and overheadpressures in shallow reactors. Thus, the oxidation reaction can becarried out at lower temperatures than would be required in a shallowervessel. When bubble column reactor 20 is used for the partial oxidationof para-xylene to crude terephthalic acid (CTA), the ability to operateat lower reaction temperatures with the same or better oxygen masstransfer rates has a number of advantages. For example, low temperatureoxidation of para-xylene reduces the amount of solvent burned during thereaction. As discussed in further detail below, low temperatureoxidation also favors the formation of small, high surface area, looselybound, easily dissolved CTA particles, which can be subjected to moreeconomical purification techniques than the large, low surface area,dense CTA particles produced by conventional high temperature oxidationprocesses.

During oxidation in reactor 20, it is preferred for the time-averagedand volume-averaged temperature of reaction medium 36 to be maintainedin the range of from about 125 to about 200° C., more preferably in therange of from about 140 to about 180° C., and most preferably in therange of from 150 to 170° C. The overhead pressure above reaction medium36 is preferably maintained in the range of from about 1 to about 20 bargauge (barg), more preferably in the range of from about 2 to about 12barg, and most preferably in the range of from 4 to 8 barg. Preferably,the pressure difference between the top of reaction medium 36 and thebottom of reaction medium 36 is in the range of from about 0.4 to about5 bar, more preferably the pressure difference is in the range of fromabout 0.7 to about 3 bars, and most preferably the pressure differenceis 1 to 2 bar. Although it is generally preferred for the overheadpressure above reaction medium 36 to be maintained at a relativelyconstant value, one embodiment of the present invention contemplatespulsing the overhead pressure to facilitate improved mixing and/or masstransfer in reaction medium 36. When the overhead pressure is pulsed, itis preferred for the pulsed pressures to range between about 60 to about140 percent of the steady-state overhead pressure recited herein, morepreferably between about 85 and about 115 percent of the steady-stateoverhead pressure recited herein, and most preferably between 95 and 105percent of the steady-state overhead pressure recited herein.

A further advantage of the high L:D ratio of reaction zone 28 is that itcan contribute to an increase in the average superficial velocity ofreaction medium 36. The term “superficial velocity” and “superficial gasvelocity”, as used herein with reference to reaction medium 36, shalldenote the volumetric flow rate of the gas phase of reaction medium 36at an elevation in the reactor divided by the horizontal cross-sectionalarea of the reactor at that elevation. The increased superficialvelocity provided by the high L:D ratio of reaction zone 28 can promotelocal mixing and increase the gas hold-up of reaction medium 36. Thetime-averaged superficial velocities of reaction medium 36 atone-quarter height, half height, and/or three-quarter height of reactionmedium 36 are preferably greater than about 0.3 meters per second, morepreferably in the range of from about 0.8 to about 5 meters per second,still more preferably in the range of from about 0.9 to about 4 metersper second, and most preferably in the range of from 1 to 3 meters persecond.

Referring again to FIG. 1, disengagement section 26 of bubble columnreactor 20 is simply a widened portion of vessel shell 22 locatedimmediately above reaction section 24. Disengagement section 26 reducesthe velocity of the upwardly-flowing gas phase in bubble column reactor20 as the gas phase rises above the upper surface 44 of reaction medium36 and approaches gas outlet 40. This reduction in the upward velocityof the gas phase helps facilitate removal of entrained liquids and/orsolids in the upwardly flowing gas phase and thereby reduces undesirableloss of certain components present in the liquid phase of reactionmedium 36.

Disengagement section 26 preferably includes a generally frustoconicaltransition wall 54, a generally cylindrical broad sidewall 56, and anupper head 58. The narrow lower end of transition wall 54 is coupled tothe top of cylindrical main body 46 of reaction section 24. The wideupper end of transition wall 54 is coupled to the bottom of broadsidewall 56. It is preferred for transition wall 54 to extend upwardlyand outwardly from its narrow lower end at an angle in the range of fromabout 10 to about 70 degrees from vertical, more preferably in the rangeof about 15 to about 50 degrees from vertical, and most preferably inthe range of from 15 to 45 degrees from vertical. Broad sidewall 56 hasa maximum diameter “X” that is generally greater than the maximumdiameter “D” of reaction section 24, though when the upper portion ofreaction section 24 has a smaller diameter than the overall maximumdiameter of reaction section 24, then X may actually be smaller than D.In a preferred embodiment of the present invention, the ratio of thediameter of broad sidewall 56 to the maximum diameter of reactionsection 24 “X:D” is in the range of from about 0.8:1 to about 4:1, mostpreferably in the range of from 1.1:1 to 2:1. Upper head 58 is coupledto the top of broad sidewall 56. Upper head 58 is preferably a generallyelliptical head member defining a central opening that permits gas toescape disengagement zone 30 via gas outlet 40. Alternatively, upperhead 58 may be of any shape, including conical. Disengagement zone 30has a maximum height “Y” measured from the top 50 of reaction zone 28 tothe upper most portion of disengagement zone 30. The ratio of the lengthof reaction zone 28 to the height of disengagement zone 30 “L:Y” ispreferably in the range of from about 2:1 to about 24:1, more preferablyin the range of from about 3:1 to about 20:1, and most preferably in therange of from 4:1 to 16:1.

Referring now to FIGS. 1-5, the location and configuration of oxidantsparger 34 will now be discussed in greater detail. FIGS. 2 and 3 showthat oxidant sparger 34 can include a ring member 60, a cross-member 62,and a pair of oxidant entry conduits 64 a,b. Conveniently, these oxidantentry conduits 64 a,b can enter the vessel at an elevation above thering member 60 and then turn downwards as shown in FIGS. 2 and 3.Alternatively, an oxidant entry conduit 64 a,b may enter the vesselbelow the ring member 60 or on about the same horizontal plane as ringmember 60. Each oxidant entry conduit 64 a,b includes a first endcoupled to a respective oxidant inlet 66a,b formed in the vessel shell22 and a second end fluidly coupled to ring member 60. Ring member 60 ispreferably formed of conduits, more preferably of a plurality ofstraight conduit sections, and most preferably a plurality of straightpipe sections, rigidly coupled to one another to form a tubularpolygonal ring. Preferably, ring member 60 is formed of at least 3straight pipe sections, more preferably 6 to 10 pipe sections, and mostpreferably 8 pipe sections. Accordingly, when ring member 60 is formedof 8 pipe sections, it has a generally octagonal configuration.Cross-member 62 is preferably formed of a substantially straight pipesection that is fluidly coupled to and extends diagonally betweenopposite pipe sections of ring member 60. The pipe section used forcross-member 62 preferably has substantially the same diameter as thepipe sections used to form ring member 60. It is preferred for the pipesections that make up oxidant entry conduits 64 a,b, ring member 60, andcross-member 62 to have a nominal diameter greater than about 0.1 meter,more preferable in the range of from about 0.2 to about 2 meters, andmost preferably in the range of from 0.25 to 1 meters. As perhaps bestillustrated in FIG. 3, ring member 60 and cross-member 62 each present aplurality of upper oxidant openings 68 for discharging the oxidantstream upwardly into reaction zone 28. As perhaps best illustrated inFIG. 4, ring member 60 and/or cross-member 62 can present one or morelower oxidant openings 70 for discharging the oxidant stream downwardlyinto reaction zone 28. Lower oxidant openings 70 can also be used todischarge liquids and/or solids that might intrude within ring member 60and/or cross-member 62. In order to prevent solids from building upinside oxidant sparger 34, a liquid stream can be continuously orperiodically passed through sparger 34 to flush out any accumulatedsolids.

Referring again to FIGS. 1-4, during oxidation in bubble column reactor20, oxidant streams are forced through oxidant inlets 66 a,b and intooxidant entry conduits 64 a,b, respectively. The oxidant streams arethen transported via oxidant entry conduits 64 a,b to ring member 60.Once the oxidant stream has entered ring member 60, the oxidant streamis distributed throughout the internal volumes of ring member 60 andcross-member 62. The oxidant stream is then forced out of oxidantsparger 34 and into reaction zone 28 via upper and lower oxidantopenings 68,70 of ring member 60 and cross-member 62.

The outlets of upper oxidant openings 68 are laterally spaced from oneanother and are positioned at substantially the same elevation inreaction zone 28. Thus, the outlets of upper oxidant openings 68 aregenerally located along a substantially horizontal plane defined by thetop of oxidant sparger 34. The outlets of lower oxidant openings 70 arelaterally spaced from one another and are positioned at substantiallythe same elevation in reaction zone 28. Thus, the outlets of loweroxidant openings 70 are generally located along a substantiallyhorizontal plane defined by the bottom of oxidant sparger 34.

In one embodiment of the present invention, oxidant sparger 34 has atleast about 20 upper oxidant openings 68 formed therein. Morepreferably, oxidant sparger 34 has in the range of from about 40 toabout 800 upper oxidant openings formed therein. Most preferably,oxidant sparger 34 has in the range of from 60 to 400 upper oxidantopenings 68 formed therein. Oxidant sparger 34 preferably has at leastabout 1 lower oxidant opening 70 formed therein. More preferably,oxidant sparger 34 has in the range of from about 2 to about 40 loweroxidant openings 70 formed therein. Most preferably, oxidant sparger 34has in the range of from 8 to 20 lower oxidant openings 70 formedtherein. The ratio of the number of upper oxidant openings 68 to loweroxidant openings 70 in oxidant sparger 34 is preferably in the range offrom about 2:1 to about 100:1, more preferably in the range of fromabout 5:1 to about 25:1, and most preferably in the range of from 8:1 to15:1. The diameters of substantially all upper and lower oxidantopenings 68,70 are preferably substantially the same, so that the ratioof the volumetric flow rate of the oxidant stream out of upper and loweropenings 68,70 is substantially the same as the ratios, given above, forthe relative number of upper and lower oxidant openings 68,70.

FIG. 5 illustrates the direction of oxidant discharge from upper andlower oxidant openings 68,70. With reference to upper oxidant openings68, it is preferred for at least a portion of upper oxidant openings 68to discharge the oxidant stream in at an angle “A” that is skewed fromvertical. It is preferred for the percentage of upper oxidant openings68 that are skewed from vertical by angle “A” to be in the range of fromabout 30 to about 90 percent, more preferably in the range of from about50 to about 80 percent, still more preferably in the range of from 60 to75 percent, and most preferably about 67 percent. The angle “A” ispreferably in the range of from about 5 to about 60 degrees, morepreferably in the range of from about 10 to about 45 degrees, and mostpreferably in the range of from 15 to 30 degrees. As for lower oxidantopenings 70, it is preferred that substantially all of lower oxidantopenings 70 are located near the bottom-most portion of the ring member60 and/or cross-member 62. Thus, any liquids and/or solids that may haveunintentionally entered oxidant sparger 34 can be readily dischargedfrom oxidant sparger 34 via lower oxidant openings 70. Preferably, loweroxidant openings 70 discharge the oxidant stream downwardly at asubstantially vertical angle. For purposes of this description, an upperoxidant opening can be any opening that discharges an oxidant stream ina generally upward direction (i.e., at an angle above horizontal), and alower oxidant opening can be any opening that discharges an oxidantstream in a generally downward direction (i.e., at an angle belowhorizontal).

In many conventional bubble column reactors containing a multi-phasereaction medium, substantially all of the reaction medium located belowthe oxidant sparger (or other mechanism for introducing the oxidantstream into the reaction zone) has a very low gas hold-up value. Asknown in the art, “gas hold-up” is simply the volume fraction of amulti-phase medium that is in the gaseous state. Zones of low gashold-up in a medium can also be referred to as “unaerated” zones. Inmany conventional slurry bubble column reactors, a significant portionof the total volume of the reaction medium is located below the oxidantsparger (or other mechanism for introducing the oxidant stream into thereaction zone). Thus, a significant portion of the reaction mediumpresent at the bottom of conventional bubble column reactors isunaerated.

It has been discovered that minimizing the amount of unaerated zones ina reaction medium subjected to oxidization in a bubble column reactorcan minimize the generation of certain types of undesirable impurities.Unaerated zones of a reaction medium contain relatively few oxidantbubbles. This low volume of oxidant bubbles reduces the amount ofmolecular oxygen available for dissolution into the liquid phase of thereaction medium. Thus, the liquid phase in an unaerated zone of thereaction medium has a relatively low concentration of molecular oxygen.These oxygen-starved, unaerated zones of the reaction medium have atendency to promote undesirable side reactions, rather than the desiredoxidation reaction. For example, when para-xylene is partially oxidizedto form terephthalic acid, insufficient oxygen availability in theliquid phase of the reaction medium can cause the formation ofundesirably high quantities of benzoic acid and coupled aromatic rings,notably including highly undesirable colored molecules known asfluorenones and anthraquinones.

In accordance with one embodiment of the present invention, liquid-phaseoxidation is carried out in a bubble column reactor configured andoperated in a manner such that the volume fraction of the reactionmedium with low gas hold-up values is minimized. This minimization ofunaerated zones can be quantified by theoretically partitioning theentire volume of the reaction medium into 2,000 discrete horizontalslices of uniform volume. With the exception of the highest and lowesthorizontal slices, each horizontal slice is a discrete volume bounded onits sides by the sidewall of the reactor and bounded on its top andbottom by imaginary horizontal planes. The highest horizontal slice isbounded on its bottom by an imaginary horizontal plane and on its top bythe upper surface of the reaction medium. The lowest horizontal slice isbounded on its top by an imaginary horizontal plane and on its bottom bythe lower end of the vessel. Once the reaction medium has beentheoretically partitioned into 2,000 discrete horizontal slices of equalvolume, the time-averaged and volume-averaged gas hold-up of eachhorizontal slice can be determined. When this method of quantifying theamount of unaerated zones is employed, it is preferred for the number ofhorizontal slices having a time-averaged and volume-averaged gas hold-upless than 0.1 to be less than 30, more preferably less than 15, stillmore preferably less than 6, even more preferably less than 4, and mostpreferably less than 2. It is preferred for the number of horizontalslices having a gas hold-up less than 0.2 to be less than 80, morepreferably less than 40, still more preferably less than 20, even morepreferably less than 12, and most preferably less than 5. It ispreferred for the number of horizontal slices having a gas hold-up lessthan 0.3 to be less than 120, more preferably less than 80, still morepreferably less than 40, even more preferably less than 20, and mostpreferably less than 15.

Referring again to FIGS. 1 and 2, it has been discovered thatpositioning oxidant sparger 34 lower in reaction zone 28 providesseveral advantages, including reduction of the amount of unaerated zonesin reaction medium 36. Given a height “H” of reaction medium 36, alength “L” of reaction zone 28, and a maximum diameter “D” of reactionzone 28, it is preferred for a majority (i.e., >50 percent by weight) ofthe oxidant stream to be introduced into reaction zone 28 within about0.025 H, 0.022 L, and/or 0.25 D of lower end 52 of reaction zone 28.More preferably, a majority of the oxidant stream is introduced intoreaction zone 28 within about 0.02 H, 0.018 L, and/or 0.2 D of lower end52 of reaction zone 28. Most preferably, a majority of the oxidantstream is introduced into reaction zone 28 within 0.015 H, 0.013 L,and/or 0.15 D of lower end 52 of reaction zone 28.

In the embodiment illustrated in FIG. 2, the vertical distance “Y₁”between lower end 52 of reaction zone 28 and the outlet of upper oxidantopenings 68 of oxidant sparger 34 is less than about 0.25 H, 0.022 L,and/or 0.25 D, so that substantially all of the oxidant stream entersreaction zone 28 within about 0.25 H, 0.022 L, and/or 0.25 D of lowerend 52 of reaction zone 28. More preferably, Y₁ is less than about 0.02H, 0.018 L, and/or 0.2 D. Most preferably, Y₁ is less than 0.015 H,0.013 L, and/or 0.15 D, but more than 0.005 H, 0.004 L, and/or 0.06 D.FIG. 2 illustrates a tangent line 72 at the location where the bottomedge of cylindrical main body 46 of vessel shell 22 joins with the topedge of elliptical lower head 48 of vessel shell 22. Alternatively,lower head 48 can be of any shape, including conical, and the tangentline is still defined as the bottom edge of cylindrical main body 46.The vertical distance “Y₂” between tangent line 72 and the top ofoxidant sparger 34 is preferably at least about 0.0012 H, 0.001 L,and/or 0.01 D; more preferably at least about 0.005 H, 0.004 L, and/or0.05 D; and most preferably at least 0.01 H, 0.008 L, and/or 0.1 D. Thevertical distance “Y₃” between lower end 52 of reaction zone 28 and theoutlet of lower oxidant openings 70 of oxidant sparger 34 is preferablyless than about 0.015 H, 0.013 L, and/or 0.15 D; more preferably lessthan about 0.012 H, 0.01 L, and/or 0.1 D; and most preferably less than0.01 H, 0.008 L, and/or 0.075 D, but more than 0.003 H, 0.002 L, and/or0.025 D.

In a preferred embodiment of the present invention, the openings thatdischarge the oxidant stream and the feed stream into the reaction zoneare configured so that the amount (by weight) of the oxidant or feedstream discharged from an opening is directly proportional to the openarea of the opening. Thus, for example, if 50 percent of the cumulativeopen area defined by all oxidants openings is located within 0.15 D ofthe bottom of the reaction zone, then 50 weight percent of the oxidantstream enters the reaction zone within 0.15 D of the bottom of thereaction zone and vice-versa.

In addition to the advantages provided by minimizing unaerated zones(i.e., zones with low gas hold-up) in reaction medium 36, it has beendiscovered that oxidation can be enhanced by maximizing the gas hold-upof the entire reaction medium 36. Reaction medium 36 preferably hastime-averaged and volume-averaged gas hold-up of at least about 0.4,more preferably in the range of from about 0.6 to about 0.9, and mostpreferably in the range of from 0.65 to 0.85. Several physical andoperational attributes of bubble column reactor 20 contribute to thehigh gas hold-up discussed above. For example, for a given reactor sizeand flow of oxidant stream, the high L:D ratio of reaction zone 28yields a lower diameter which increases the superficial velocity inreaction medium 36 which in turn increases gas hold-up. Additionally,the actual diameter of a bubble column and the L:D ratio are known toinfluence the average gas hold-up even for a given constant superficialvelocity. In addition, the minimization of unaerated zones, particularlyin the bottom of reaction zone 28, contributes to an increased gashold-up value. Further, the overhead pressure and mechanicalconfiguration of the bubble column reactor can affect operatingstability at the high superficial velocities and gas hold-up valuesdisclosed herein.

Furthermore, the inventors have discovered the importance of operatingwith an optimized overhead pressure to obtain increased gas hold-up andincreased mass transfer. It might seem that operating with a loweroverhead pressure, which reduces the solubility of molecular oxygenaccording to a Henry's Law effect, would reduce the mass transfer rateof molecular oxygen from gas to liquid. In a mechanically agitatedvessel, such is typically the case because aeration levels and masstransfer rates are dominated by agitator design and overhead pressure.However, in a bubble column reactor according to a preferred embodimentof the present invention, it has been discovered how to use a loweroverhead pressure to cause a given mass of gas-phase oxidant stream tooccupy more volume, increasing the superficial velocity in reactionmedium 36 and in turn increasing the gas hold-up and transfer rate ofmolecular oxygen.

The balance between bubble coalescence and breakup is an extremelycomplicated phenomenon, leading on the one hand to a tendency to foam,which reduces internal circulation rates of the liquid phase and whichmay require very, very large disengaging zones, and on the other hand toa tendency to fewer, very large bubbles that give a lower gas hold-upand lower mass transfer rate from the oxidant stream to the liquidphase. Concerning the liquid phase, its composition, density, viscosityand surface tension, among other factors, are known to interact in avery complicated manner to produce very complicated results even in theabsence of a solid-phase. For example, laboratory investigators havefound it useful to qualify whether “water” is tap water, distilledwater, or de-ionized water, when reporting and evaluating observationsfor even simple water-air bubble columns. For complex mixtures in theliquid phase and for the addition of a solid phase, the degree ofcomplexity rises further. The surface irregularities of individualparticles of solids, the average size of solids, the particle sizedistribution, the amount of solids relative to the liquid phase, and theability of the liquid to wet the surface of the solid, among otherthings, are all important in their interaction with the liquid phase andthe oxidant stream in establishing what bubbling behavior and naturalconvection flow patterns will result.

Thus, the ability of the bubble column reactor to function usefully withthe high superficial velocities and high gas hold-up disclosed hereindepends, for example, on an appropriate selection of: (1) thecomposition of the liquid phase of the reaction medium; (2) the amountand type of precipitated solids, both of which can be adjusted byreaction conditions; (3) the amount of oxidant stream fed to thereactor; (4) the overhead pressure, which affects the volumetric flow ofoxidant stream, the stability of bubbles, and, via the energy balance,the reaction temperature; (5) the reaction temperature itself, whichaffects the fluid properties, the properties of precipitated solids, andthe specific volume of the oxidant stream; and (6) the geometry andmechanical details of the reaction vessel, including the L:D ratio.

Referring again to FIG. 1, it has been discovered that improveddistribution of the oxidizable compound (e.g., para-xylene) in reactionmedium 36 can be provided by introducing the liquid-phase feed streaminto reaction zone 28 at multiple vertically-spaced locations.Preferably, the liquid-phase feed stream is introduced into reactionzone 28 via at least 3 feed openings, more preferably at least 4 feedopenings. As used herein, the term “feed openings” shall denote openingswhere the liquid-phase feed stream is discharged into reaction zone 28for mixing with reaction medium 36. It is preferred for at least 2 ofthe feed openings to be vertically-spaced from one another by at leastabout 0.5 D, more preferably at least about 1.5 D, and most preferablyat least 3 D. However, it is preferred for the highest feed opening tobe vertically-spaced from the lowest oxidant opening by not more thanabout 0.75 H, 0.65 L, and/or 8 D; more preferably not more than about0.5 H, 0.4 L, and/or 5 D; and most preferably not more than 0.4 H, 0.35L, and/or 4 D.

Although it is desirable to introduce the liquid-phase feed stream atmultiple vertical locations, it has also been discovered that improveddistribution of the oxidizable compound in reaction medium 36 isprovided if the majority of the liquid-phase feed stream is introducedinto the bottom half of reaction medium 36 and/or reaction zone 28.Preferably, at least about 75 weight percent of the liquid-phase feedstream is introduced into the bottom half of reaction medium 36 and/orreaction zone 28. Most preferably, at least 90 weight percent of theliquid-phase feed stream is introduced into the bottom half of reactionmedium 36 and/or reaction zone 28. In addition, it is preferred for atleast about 30 weight percent of the liquid-phase feed stream to beintroduced into reaction zone 28 within about 1.5 D of the lowestvertical location where the oxidant stream is introduced into reactionzone 28. This lowest vertical location where the oxidant stream isintroduced into reaction zone 28 is typically at the bottom of oxidantsparger; however, a variety of alternative configurations forintroducing the oxidant stream into reaction zone 28 are contemplated bya preferred embodiment of the present invention. Preferably, at leastabout 50 weight percent of the liquid-phase feed is introduced withinabout 2.5 D of the lowest vertical location where the oxidant stream isintroduced into reaction zone 28. Preferably, at least about 75 weightpercent of the liquid-phase feed stream is introduced within about 5 Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28.

Each feed opening defines an open area through which the feed isdischarged. It is preferred that at least about 30 percent of thecumulative open area of all the feed inlets is located within about 1.5D of the lowest vertical location where the oxidant stream is introducedinto reaction zone 28. Preferably, at least about 50 percent of thecumulative open area of all the feed inlets is located within about 2.5D of the lowest vertical location where the oxidant stream is introducedinto reaction zone 28. Preferably, at least about 75 percent of thecumulative open area of all the feed inlets is located within about 5 Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28.

Referring again to FIG. 1, in one embodiment of the present invention,feed inlets 32 a,b,c,d are simply a series of vertically-alignedopenings along one side of vessel shell 22. These feed openingspreferably have substantially similar diameters of less than about 7centimeters, more preferably in the range of from about 0.25 to about 5centimeters, and most preferably in the range of from 0.4 to 2centimeters. Bubble column reactor 20 is preferably equipped with asystem for controlling the flow rate of the liquid-phase feed stream outof each feed opening. Such flow control system preferably includes anindividual flow control valve 74 a,b,c,d for each respective feed inlet32 a,b,c,d. In addition, it is preferred for bubble column reactor 20 tobe equipped with a flow control system that allows at least a portion ofthe liquid-phase feed stream to be introduced into reaction zone 28 atan elevated inlet superficial velocity of at least about 2 meters persecond, more preferably at least about 5 meters per second, still morepreferably at least about 6 meters per second, and most preferably inthe range of from 8 to 20 meters per second. As used herein, the term“inlet superficial velocity” denotes the time-averaged volumetric flowrate of the feed stream out of the feed opening divided by the area ofthe feed opening. Preferably, at least about 50 weight percent of thefeed stream is introduced into reaction zone 28 at an elevated inletsuperficial velocity. Most preferably, substantially all the feed streamis introduced into reaction zone 28 at an elevated inlet superficialvelocity.

Referring now to FIGS. 6-7, an alternative system for introducing theliquid-phase feed stream into reaction zone 28 is illustrated. In thisembodiment, the feed stream is introduced into reaction zone 28 at fourdifferent elevations. Each elevation is equipped with a respective feeddistribution system 76 a,b,c,d. Each feed distribution system 76includes a main feed conduit 78 and a manifold 80. Each manifold 80 isprovided with at least two outlets 82,84 coupled to respective insertconduits 86,88, which extend into reaction zone 28 of vessel shell 22.Each insert conduit 86,88 presents a respective feed opening 87,89 fordischarging the feed stream into reaction zone 28. Feed openings 87,89preferably have substantially similar diameters of less than about 7centimeters, more preferably in the range of from about 0.25 to about 5centimeters, and most preferably in the range of from 0.4 to 2centimeters. It is preferred for feed openings 87,89 of each feeddistribution system 76 a,b,c,d to be diametrically opposed so as tointroduce the feed stream into reaction zone 28 in opposite directions.Further, it is preferred for the diametrically opposed feed openings86,88 of adjacent feed distribution systems 76 to be oriented at 90degrees of rotation relative to one another. In operation, theliquid-phase feed stream is charged to main feed conduit 78 andsubsequently enters manifold 80. Manifold 80 distributes the feed streamevenly for simultaneous introduction on opposite sides of reactor 20 viafeed openings 87,89.

FIG. 8 illustrates an alternative configuration wherein each feeddistribution system 76 is equipped with bayonet tubes 90,92 rather thaninsert conduits 86,88 (shown in FIG. 7). Bayonet tubes 90,92 projectinto reaction zone 28 and include a plurality of small feed openings94,96 for discharging the liquid-phase feed into reaction zone 28. It ispreferred for the small feed openings 94,96 of bayonet tubes 90,92 tohave substantially the same diameters of less than about 50 millimeters,more preferably about 2 to about 25 millimeters, and most preferably 4to 15 millimeters.

FIGS. 9-11 illustrate an alternative feed distribution system 100. Feeddistribution system 100 introduces the liquid-phase feed stream at aplurality of vertically-spaced and laterally-spaced locations withoutrequiring multiple penetrations of the sidewall of bubble column reactor20. Feed introduction system 100 generally includes a single inletconduit 102, a header 104, a plurality of upright distribution tubes106, a lateral support mechanism 108, and a vertical support mechanism110. Inlet conduit 102 penetrates the sidewall of main body 46 of vesselshell 22. Inlet conduit 102 is fluidly coupled to header 104. Header 104distributes the feed stream received from inlet conduit 102 evenly amongupright distribution tubes 106. Each distribution tube 106 has aplurality of vertically-spaced feed openings 112 a,b,c,d for dischargingthe feed stream into reaction zone 28. Lateral support mechanism 108 iscoupled to each distribution tube 106 and inhibits relative lateralmovement of distribution tubes 106. Vertical support mechanism 110 ispreferably coupled to lateral support mechanism 108 and to the top ofoxidant sparger 34. Vertical support mechanism 110 substantiallyinhibits vertical movement of distribution tubes 106 in reaction zone28. It is preferred for feed openings 112 to have substantially the samediameters of less than about 50 millimeters, more preferably about 2 toabout 25 millimeters, and most preferably 4 to 15 millimeters. Thevertical spacing of feed openings 112 of feed distribution system 100illustrated in FIGS. 9-11 can be substantially the same as describedabove with reference to the feed distribution system of FIG. 1.

It has been discovered that the flow patterns of the reaction medium inmany bubble column reactors can permit uneven azimuthal distribution ofthe oxidizable compound in the reaction medium, especially when theoxidizable compound is primarily introduced along one side of thereaction medium. As used herein, the term “azimuthal” shall denote anangle or spacing around the upright axis of elongation of the reactionzone. As used herein, “upright” shall mean within 45° of vertical. Inone embodiment of the present invention, the feed stream containing theoxidizable compound (e.g., para-xylene) is introduced into the reactionzone via a plurality of azimuthally-spaced feed openings. Theseazimuthally-spaced feed openings can help prevent regions of excessivelyhigh and excessively low oxidizable compound concentrations in thereaction medium. The various feed introduction systems illustrated inFIGS. 6-11 are examples of systems that provide proper azimuthal spacingof feed openings.

Referring again to FIG. 7, in order to quantify the azimuthally-spacedintroduction of the liquid-phase feed stream into the reaction medium,the reaction medium can be theoretically partitioned into four uprightazimuthal quadrants “Q₁,Q₂,Q₃,Q₄” of approximately equal volume. Theseazimuthal quadrants “Q₁,Q₂,Q₃,Q₄” are defined by a pair of imaginaryintersecting perpendicular vertical planes “P₁,P₂” extending beyond themaximum vertical dimension and maximum radial dimension of the reactionmedium. When the reaction medium is contained in a cylindrical vessel,the line of intersection of the imaginary intersecting vertical planesP₁,P₂ will be approximately coincident with the vertical centerline ofthe cylinder, and each azimuthal quadrant Q₁,Q₂,Q₃,Q₄ will be agenerally wedge-shaped vertical volume having a height equal to theheight of the reaction medium. It is preferred for a substantial portionof the oxidizable compound to be discharged into the reaction medium viafeed openings located in at least two different azimuthal quadrants.

In a preferred embodiment of the present invention, not more than about80 weight percent of the oxidizable compound is discharged into thereaction medium through feed openings that can be located in a singleazimuthal quadrant. More preferably, not more than about 60 weightpercent of the oxidizable compound is discharged into the reactionmedium through feed openings that can be located in a single azimuthalquadrant. Most preferably, not more than 40 weight percent of theoxidizable compound is discharged into the reaction medium through feedopenings that can be located in a single azimuthal quadrant. Theseparameters for azimuthal distribution of the oxidizable compound aremeasured when the azimuthal quadrants are azimuthally oriented such thatthe maximum possible amount of oxidizable compound is being dischargedinto one of the azimuthal quadrants. For example, if the entire feedstream is discharged into the reaction medium via two feed openings thatare azimuthally spaced from one another by 89 degrees, for purposes ofdetermining azimuthal distribution in four azimuthal quadrants, 100weight percent of the feed stream is discharged into the reaction mediumin a single azimuthal quadrant because the azimuthal quadrants can beazimuthally oriented in such a manner that both of the feed openings arelocated in a single azimuthal quadrant.

In addition to the advantages associated with the properazimuthal-spacing of the feed openings, it has also been discovered thatproper radial spacing of the feed openings in a bubble column reactorcan also be important. It is preferred for a substantial portion of theoxidizable compound introduced into the reaction medium to be dischargedvia feed openings that are radially spaced inwardly from the sidewall ofthe vessel. Thus, in one embodiment of the present invention, asubstantial portion of the oxidizable compound enters the reaction zonevia feed openings located in a “preferred radial feed zone” that isspaced inwardly from the upright sidewalls defining the reaction zone.

Referring again to FIG. 7, the preferred radial feed zone “FZ” can takethe shape of a theoretical upright cylinder centered in reaction zone 28and having an outer diameter “D_(O)” of 0.9 D, where “D” is the diameterof reaction zone 28. Thus, an outer annulus “OA” having a thickness of0.05 D is defined between the preferred radial feed zone FZ and theinside of the sidewall defining reaction zone 28. It is preferred forlittle or none of the oxidizable compound to be introduced into reactionzone 28 via feed openings located in this outer annulus OA.

In another embodiment, it is preferred for little or none of theoxidizable compound to be introduced into the center of reaction zone28. Thus, as illustrated in FIG. 8, the preferred radial feed zone FZcan take the shape of a theoretical upright annulus centered in reactionzone 28, having an outer diameter D_(O) of 0.9 D, and having an innerdiameter D_(I) of 0.2 D. Thus, in this embodiment, an inner cylinder IChaving a diameter of 0.2 D is “cut out” of the center of the preferredradial feed zone FZ. It is preferred for little or none of theoxidizable compound to be introduced into reaction zone 28 via feedopenings located in this inner cylinder IC.

In a preferred embodiment of the present invention, a substantialportion of the oxidizable compound is introduced into reaction medium 36via feed openings located in the preferred radial feed zone, regardlessof whether the preferred radial feed zone has the cylindrical or annularshape described above. More preferably, at least about 25 weight percentof the oxidizable compound is discharged into reaction medium 36 viafeed openings located in the preferred radial feed zone. Still morepreferably, at least about 50 weight percent of the oxidizable compoundis discharged into reaction medium 36 via feed openings located in thepreferred radial feed zone. Most preferably, at least 75 weight percentof the oxidizable compound is discharged into reaction medium 36 viafeed openings located in the preferred radial feed zone.

Although the theoretical azimuthal quadrants and theoretical preferredradial feed zone illustrated in FIGS. 7 and 8 are described withreference to the distribution of the liquid-phase feed stream, it hasbeen discovered that proper azimuthal and radial distribution of thegas-phase oxidant stream can also provide certain advantages. Thus, inone embodiment of the present invention, the description of theazimuthal and radial distribution of the liquid-phase feed stream,provided above, also applies to the manner in which the gas-phaseoxidant stream is introduced into the reaction medium 36.

Referring now to FIGS. 12-15, an alternative oxidant sparger 200 isillustrated as generally comprising a ring member 202 and a pair ofoxidant entry conduits 204,206. Oxidant sparger 200 of FIGS. 12-15 issimilar to oxidant sparger 34 of FIGS. 1-11 with the following threeprimary differences: (1) oxidant sparger 200 does not include a diagonalcross-member; (2) the upper portion of ring member 202 has no openingsfor discharging the oxidant in an upward direction; and (3) oxidantsparger 200 has many more openings in the lower portion of ring member202.

As perhaps best illustrated in FIGS. 14 and 15, the bottom portion ofoxidant sparger ring 202 presents a plurality of oxidant openings 208.Oxidant openings 208 are preferably configured such that at least about1 percent of the total open area defined by oxidant openings 208 islocated below the centerline 210 (FIG. 15) of ring member 202, wherecenterline 210 is located at the elevation of the volumetric centroid ofring member 202. More preferably, at least about 5 percent of the totalopen area defined by all oxidant openings 208 is located belowcenterline 210, with at least about 2 percent of the total open areabeing defined by openings 208 that discharge the oxidant stream in agenerally downward direction within about 30 degrees of vertical. Stillmore preferably, at least about 20 percent of the total open areadefined by all oxidant openings 208 is located below centerline 210,with at least about 10 percent of the total open area being defined byopenings 208 that discharge the oxidant stream in a generally downwarddirection within 30 degrees of vertical. Most preferably, at least about75 percent of the total open area defined by all oxidant openings 208 islocated below centerline 210, with at least about 40 percent of thetotal open area being defined by openings 208 that discharge the oxidantstream in a generally downward direction within 30 degrees of vertical.The fraction of the total open area defined by all oxidant openings 208that are located above centerline 210 is preferably less than about 75percent, more preferably less than about 50 percent, still morepreferably less than about 25 percent, and most preferably less than 5percent.

As illustrated in FIGS. 14 and 15, oxidant openings 208 include downwardopenings 208 a and skewed openings 208 b. Downward openings 208 a areconfigured to discharge the oxidant stream generally downwardly at anangle within about 30 degrees of vertical, more preferably within about15 degrees of vertical, and most preferably within 5 degrees ofvertical. Skewed openings 208 b are configured to discharge the oxidantstream generally outwardly and downwardly at an angle “A” that is in therange of from about 15 to about 75 degrees from vertical, morepreferably angle A is in the range of from about 30 to about 60 degreesfrom vertical, and most preferably angle A is in the range of from 40 to50 degrees from vertical.

It is preferred for substantially all oxidant openings 208 to haveapproximately the same diameter. The diameter of oxidant openings 208 ispreferably in the range of from about 2 to about 300 millimeters, morepreferably in the range of from about 4 to about 120 millimeters, andmost preferably in the range of from 8 to 60 millimeters. The totalnumber of oxidant openings 208 in ring member 202 is selected to meetthe low pressure drop criteria detailed below. Preferably, the totalnumber of oxidant openings 208 formed in ring member 202 is at leastabout 10, more preferably the total number of oxidant openings 208 is inthe range of from about 20 to about 200, and most preferably the totalnumber of oxidant openings 208 is in the range of from 40 to 100.

Although FIGS. 12-15 illustrate a very specific configuration foroxidant sparger 200, it is now noted that a variety of oxidant spargerconfigurations can be employed to achieve the advantages describedherein. For example, the oxidant sparger does not necessarily need tohave the octagonal ring member configuration illustrated in FIGS. 12-13.Rather, it is possible for the oxidant sparger to be formed of anyconfiguration of flow conduit(s) that employs a plurality ofspaced-apart openings for discharging the oxidant stream. The size,number, and discharge direction of the oxidant openings in the flowconduit are preferably within the ranges stated above. Further, theoxidant sparger is preferably configured to provide the azimuthal andradial distribution of molecular oxygen described above.

Regardless of the specific configuration of the oxidant sparger, it ispreferred for the oxidant sparger to be physically configured andoperated in a manner that minimizes the pressure drop associated withdischarging the oxidant stream out of the flow conduit(s), through theoxidant openings, and into the reaction zone. Such pressure drop iscalculated as the time-averaged static pressure of the oxidant streaminside the flow conduit at oxidant inlets 66 a,b of the oxidant spargerminus the time-averaged static pressure in the reaction zone at theelevation where one-half of the oxidant stream is introduced above thatvertical location and one-half of the oxidant stream is introduced belowthat vertical location. In a preferred embodiment of the presentinvention, the time-averaged pressure drop associated with dischargingthe oxidant stream from the oxidant sparger is less than about 0.3megapascal (MPa), more preferably less than about 0.2 MPa, still morepreferably less than about 0.1 MPa, and most preferably less than 0.05MPa. Under the preferred operating conditions of the bubble columnreactor described herein, the pressure of the oxidant stream inside theflow conduit(s) of the oxidant sparger is preferably in the range offrom about 0.35 to about 1 MPa, more preferably in the range of fromabout 0.45 to about 0.85 MPa, and most preferably in the range of from0.5 to 0.7 MPa.

As alluded to earlier with reference to the oxidant spargerconfiguration illustrated in FIGS. 2-5, it may be desirable tocontinuously or periodically flush the oxidant sparger with a liquid(e.g., acetic acid, water, and/or para-xylene) to prevent fouling of theoxidant sparger with solids. When such a liquid flush is employed, it ispreferred for an effective amount of the liquid (i.e., not just theminor amount of liquid droplets that might naturally be present in theoxidant stream) to be passed through the oxidant sparger and out of theoxidant openings for at least one period of more than one minute eachday. When a liquid is continuously or periodically discharged from theoxidant sparger, it is preferred for the time-averaged ratio of the massflow rate of the liquid through the oxidant sparger to the mass flowrate of the molecular oxygen through the oxidant sparger to be in therange of from about 0.05:1 to about 30:1, or in the range of from about0.1:1 to about 2:1, or even in the range of from 0.2:1 to 1:1.

In one embodiment of the present invention, a significant portion of theoxidizable compound (e.g., para-xylene) can be introduced into thereaction zone through the oxidant sparger. In such a configuration, itis preferred for the oxidizable compound and the molecular oxygen to bedischarged from the oxidant sparger through the same openings in theoxidant sparger. As noted above, the oxidizable compound is typically aliquid at STP. Therefore, in this embodiment, a two-phase stream may bedischarged from the oxidant sparger, with the liquid phase comprisingthe oxidizable compound and the gas phase comprising the molecularoxygen. It should be recognized, however, that at least a portion of theoxidizable compound may be in a gaseous state when discharged from theoxidant sparger. In one embodiment, the liquid phase discharged from theoxidant sparger is formed predominately of the oxidizable compound. Inanother embodiment, the liquid phase discharged from the oxidant spargerhas substantially the same composition as the feed stream, describedabove. When the liquid phase discharged from the oxidant sparger hassubstantially the same composition as the feed stream, such liquid phasemay comprise a solvent and/or a catalyst system in the amounts andratios described above with reference to the composition of the feedstream.

In one embodiment of the present invention, it is preferred for at leastabout 10 weight percent of all the oxidizable compound introduced intothe reaction zone to be introduced via the oxidant sparger, morepreferably at least about 40 weight percent of the oxidizable compoundis introduced into the reaction zone via the oxidant sparger, and mostpreferably at least 80 weight percent of the oxidizable compound isintroduced into the reaction zone via the oxidant sparger. When all orpart of the oxidizable compound is introduced into the reaction zone viathe oxidant sparger, it is preferred for at least about 10 weightpercent of all the molecular oxygen introduced into the reaction zone tobe introduced via the same oxidant sparger, more preferably at leastabout 40 weight percent of the oxidizable compound is introduced intothe reaction zone via the same oxidant sparger, and most preferably atleast 80 weight percent of the oxidizable compound is introduced intothe reaction zone via the same oxidant sparger. When a significantportion of the oxidizable compound is introduced into the reaction zonevia the oxidant sparger, it is preferred for one or more temperaturesensing devices (e.g., thermocouples) to be disposed in the oxidantsparger. These temperature sensors can be employed to help to make surethe temperature in the oxidant sparger does not become dangerously high.

Referring now to FIGS. 16-18, bubble column reactor 20 is illustrated asincluding an internal deaeration vessel 300 disposed in the bottom ofreaction zone 28 near slurry outlet 38. It has been discovered thatimpurity-forming side reactions occur at a relatively high rate duringdeaeration of reaction medium 36. As used herein, “deaeration” shalldenote the disengagement of a gas phase from multi-phase reactionmedium. When reaction medium 36 is highly aerated (>0.3 gas hold-up),impurity formation is minimal. When reaction medium 36 is highlyunaerated (<0.01 gas hold-up), impurity formation is also minimal.However, when reaction medium is partially-aerated (0.01-0.3 gashold-up), undesirable side reactions are promoted and increasedimpurities are generated. Deaeration vessel 300 addresses this and otherproblems by minimizing the volume of reaction medium 36 in apartially-aerated stated, and by minimizing the time it takes todeaerate reaction medium 36. A substantially deaerated slurry isproduced from the bottom of deaeration vessel 300 and exits reactor 20via slurry outlet 38. The substantially deaerated slurry preferablycontains less than about 5 volume percent gas phase, more preferablyless than about 2 volume percent gas phase, and most preferably lessthan 1 volume percent gas phase.

In FIG. 16, bubble column reactor 20 is illustrated as including a levelcontroller 302 and a flow control valve 304. Level controller 302 andflow control valve 304 cooperate to maintain reaction medium 36 at asubstantially constant elevation in reaction zone 28. Level controller302 is operable to sense (e.g., by differential pressure level sensingor by nuclear level sensing) the elevation of upper surface 44 ofreaction medium 36 and generate a control signal 306 responsive to theelevation of reaction medium 36. Flow control valve 304 receives controlsignal 306 and adjusts the flow rate of a slurry through a slurry outletconduit 308. Thus, the flow rate of the slurry out of slurry outlet 38can vary between a maximum slurry volumetric flow rate (F_(max)) whenthe elevation of reaction medium 36 is too high and a minimum slurryvolumetric flow rate (F_(min)) when the elevation of reaction medium 36is too low.

In order to remove solid-phase oxidation product from reaction zone 28,a portion must first pass through deaeration vessel 300. Deaerationvessel 300 provides a low-turbulence internal volume that permits thegas phase of reaction medium 36 to naturally rise out of the liquid andsolid phases of reaction medium 36 as the liquid and solids flowdownwardly toward slurry outlet 38. The rising of the gas phase out ofthe liquid and solid phases is caused by the natural upward buoyancy ofthe gas phase in the liquid and solid phases. When deaeration vessel 300is employed, the transitioning of reaction medium 36 from afully-aerated, three-phase medium to a fully-deaerated, two-phase slurryis quick and efficient.

Referring now to FIGS. 17 and 18, deaeration vessel 300 includes agenerally upright sidewall 308 defining a deaeration zone 312therebetween. Preferably, sidewall 308 extends upwardly within about 30degrees of vertical, more preferably within about 10 degrees ofvertical. Most preferably, sidewall 308 is substantially vertical.Deaeration zone 312 is separate from reaction zone 28 and has height “h”and a diameter “d.” An upper end 310 of sidewall 308 is open so as toreceive reaction medium from reaction zone 28 into internal volume 312.The lower end of sidewall 308 is fluidly coupled to slurry outlet 38 viaa transition section 314. In certain instances, such as when the openingof slurry outlet 38 is large or when the diameter “d” of sidewall 308 issmall, transition section 314 can be eliminated. As perhaps bestillustrated in FIG. 18, deaeration vessel 300 can also include a vortexbreaker 316 disposed in deaeration zone 312. Vortex breaker 316 can beany structure operable to inhibit the formation of vortices as the solidand liquid phases flow downwardly towards slurry outlet 38.

In order to permit proper disengagement of the gas phase from the solidand liquid phases in deaeration vessel 300, the height “h” andhorizontal cross-sectional area of internal deaeration zone 312 arecarefully selected. The height “h” and horizontal cross-sectional areaof internal deaeration zone 312 should provide sufficient distance andtime so that even when the maximum amount of slurry is being withdrawn(i.e., when slurry is being withdrawn at F_(max)), substantially all ofthe gas bubble volume can rise out of the solid and liquid phases beforethe gas bubbles reach the bottom outlet of deaeration vessel 300. Thus,it is preferred for the cross-sectional area of deaeration zone 312 tobe such that the maximum downward velocity (V_(dmax)) of the liquid andsolid phases through deaeration zone 312 is substantially less than thenatural rise velocity (V_(u)) of the gas phase bubbles through theliquid and solid phases. The maximum downward velocity (V_(dmax)) of theliquid and solid phases through deaeration zone 312 occurs at themaximum slurry volumetric flow rate (F_(max)), discussed above. Thenatural rise velocity (V_(u)) of the gas bubbles through the liquid andsolid phases varies depending on the size of the bubbles; however, thenatural rise velocity (V_(u0.5)) of 0.5 centimeter diameter gas bubblesthrough the liquid and solid phases can be used as a cut-off valuebecause substantially all of the bubble volume initially in reactionmedium 36 will be greater than 0.5 centimeters. Preferably, thecross-sectional area of deaeration zone 312 is such that V_(dmax) isless than about 75 percent of V_(u0.5), more preferably V_(dmax) is lessthan about 40 percent of V_(u0.5), most preferably V_(dmax) is less than20 percent of V_(u0.5).

The downward velocity of the liquid and solid phases in deaeration zone312 of deaeration vessel 300 is calculated as the volumetric flow rateof the deaerated slurry through slurry outlet 38 divided by the minimumcross-sectional area of deaeration zone 312. The downward velocity ofthe liquid and solid phases in deaeration zone 312 of deaeration vessel300 is preferably less than about 50 centimeters per second, morepreferably less than about 30 centimeters per second, and mostpreferably less than 10 centimeters per second.

It is now noted that although upright sidewall 308 of deaeration vessel300 is illustrated as having a cylindrical configuration, sidewall 308could comprise a plurality of sidewalls that form a variety ofconfigurations (e.g., triangular, square, or oval), so long as the wallsdefines an internal volume having an appropriate volume, cross-sectionalarea, width “d”, and height “h”. In a preferred embodiment of thepresent invention, “d” is in the range of from about 0.2 to about 2meters, more preferably in the range of from about 0.3 to about 1.5meters, and most preferably in the range of from 0.4 to 1.2 meters. In apreferred embodiment of the present invention, “h” is in the range offrom about 0.3 meters to about 5 meters, more preferably in the range offrom about 0.5 to about 3 meters, and most preferably in the range offrom 0.75 to 2 meters.

In a preferred embodiment of the present invention, sidewall 308 issubstantially vertical so that the horizontal cross-sectional area ofdeaeration zone 312 is substantially constant along the entire height“h” of deaeration zone 312. Preferably, the maximum horizontalcross-sectional area of deaeration zone 312 is less than about 25percent of the maximum horizontal cross-sectional area of reaction zone28. More preferably, the maximum horizontal cross-sectional area ofdeaeration zone 312 is in the range of from about 0.1 to about 10percent of the maximum horizontal cross-sectional area of reaction zone28. Most preferably, the maximum horizontal cross-sectional area ofdeaeration zone 312 is in the range of from 0.25 to 4 percent of themaximum horizontal cross-sectional area of reaction zone 28. Preferably,the maximum horizontal cross-sectional area of deaeration zone 312 is inthe range of from about 0.02 to about 3 square meters, more preferablyin the range of from about 0.05 to about 2 square meters, and mostpreferably in the range of from 0.1 to 1.2 square meters. The volume ofdeaeration zone 312 is preferably less than about 5 percent of the totalvolume of reaction medium 36 or reaction zone 28. More preferably, thevolume of deaeration zone 312 is in the range of from about 0.01 toabout 2 percent of the total volume of reaction medium 36 or reactionzone 28. Most preferably, the volume of deaeration zone 312 is in therange of from 0.05 to about 1 percent of the total volume of reactionmedium 36 or reaction zone 28. The volume of deaeration zone 312 ispreferably less than about 2 cubic meters, more preferably in the rangeof from about 0.01 to about 1 cubic meters, and most preferably in therange of from 0.05 to 0.5 cubic meters.

Turning now to FIG. 19, bubble column reactor 20 is illustrated asincluding an external deaeration vessel 400. In this configuration,aerated reaction medium 36 is withdrawn from reaction zone 28 via anelevated opening in the side of vessel shell 22. The withdrawn aeratedmedium is transported to external deaeration vessel 400 via an outletconduit 402 for disengagement of the gas phase from the solid and liquidphases. The disengaged gas phase exits deaeration vessel 400 via conduit404, while the substantially deaerated slurry exits deaeration vessel400 via conduit 406.

In FIG. 19, outlet conduit 402 is shown as being approximately straight,horizontal, and orthogonal to vessel shell 22. This is merely oneconvenient configuration; and outlet conduit 402 may be otherwise in anyrespect, providing that it usefully connects bubble column reactor 20with external deaeration vessel 400. Turning to conduit 404, it isuseful for this conduit to connect at or near the top deaeration vessel400 in order to control safety issues relating to a stagnant gas pocketcontaining oxidizable compound and oxidant. Furthermore, conduits 402and 404 may usefully comprise means of flow isolation, such as valves.

When reaction medium 36 is withdrawn from reactor 20 via an elevatedoutlet, as shown in FIG. 19, it is preferred for bubble column reactor20 to be equipped with a lower outlet 408 near the bottom 52 of reactionzone 28. Lower outlet 408 and a lower conduit 410, coupled thereto, canbe used to de-inventory (i.e., empty) reactor 20 during shutdowns.Preferably, one or more lower outlet 408 is provided in the bottomone-third of the height of reaction medium 36, more preferably in thebottom one-fourth of reaction medium 36, and most preferably at thelowest point of reaction zone 28.

With the elevated slurry withdrawal and deaeration system shown in FIG.19, lower conduit 410 and outlet 408 are not used to withdraw slurryfrom reaction zone 28 during oxidation. It is known in the art thatsolids tend to settle by gravity forces in unaerated and otherwiseunagitated portions of the slurry, including in stagnant flow conduits.Furthermore, the settled solids (e.g., terephthalic acid) can tend tosolidify into large agglomerates by continuing precipitation and/orcrystalline reorganization. Thus, in order to avoid plugging of lowerflow conduit 410, a fraction of the deaerated slurry from the bottom ofdeaeration vessel 400 can be used to continuously or intermittentlyflush lower conduit 410 during normal operation of reactor 20. Apreferred means of providing such a slurry flush to conduit 410 is toperiodically open a valve 412 in conduit 410 and allow a fraction of thedeaerated slurry to flow through conduit 410 and into reaction zone 28via lower opening 408. Even when valve 412 is fully or partially open,only a fraction of the deaerated slurry flows through lower conduit 410and back into reaction zone 28. The remaining fraction of the deaeratedslurry not used to flush lower conduit 410 is carried via conduit 414away from reactor 20 for further downstream processing (e.g.,purification).

During normal operation of bubble column reactor 20 over a substantiallength of time (e.g., >100 hours), it is preferred for the amount ofdeaerated slurry used to flush lower conduit 410 to be less than 50percent by weight of the total deaerated slurry produced from the bottomof deaeration vessel 400, more preferably less than about 20 percent byweight, and most preferably less than 5 percent by weight. Further, itis preferred that over a substantial length of time the average massflow rate of deaerated slurry used to flush lower conduit 410 is lessthan about 4 times the average mass flow rate of the oxidizable compoundinto reaction zone 28, more preferably less than about 2 times theaverage mass flow rate of the oxidizable compound into reaction zone 28,still more preferably less than the average mass flow rate of theoxidizable compound into reaction zone 28, and most preferably less than0.5 times the average mass flow rate of the oxidizable compound intoreaction zone 28.

Referring again to FIG. 19, deaeration vessel 400 includes asubstantially upright, preferably cylindrical sidewall 416 defining adeaeration zone 418. Deaeration zone 418 has a diameter “d” and height“h.” Height “h” is measured as the vertical distance between thelocation where the aerated reaction medium enters deaeration vessel 400and the bottom of sidewall 416. The height “h”, diameter “d”, area, andvolume of deaeration zone 418 is preferably substantially the same asdescribed above with reference to deaeration zone 312 of deaerationvessel 300 illustrated in FIGS. 16-18. In addition, deaeration vessel400 includes an upper section 420 formed by extending sidewall 416 abovedeaeration zone 418. Upper section 420 of deaeration vessel 400 may beof any height, though it preferably extends upwardly to or above thelevel of reaction medium 36 in reaction zone 28. Upper section 420ensures that the gas phase has room to properly disengage from theliquid and solid phases before exiting deaeration vessel 400 via conduit404. It is now noted that although conduit 404 is illustrated asreturning the disengaged gas phase to the disengagement zone of reactor20, conduit 404 could alternatively be coupled to vessel shell 22 at anyelevation above outlet conduit 402. Optionally, conduit 404 could becoupled to gas outlet conduit 40 so that the disengaged gas phase fromdeaeration vessel 400 is combined with the removed overhead vapor streamin conduit 40 and sent downstream for further processing.

Turning now to FIG. 20, bubble column reactor 20 is illustrated asincluding a hybrid internal-external deaeration vessel 500. In thisconfiguration, a portion of reaction medium 36 is withdrawn fromreaction zone 28 through a relatively large elevated opening 502 in thesidewall of vessel shell 22. The withdrawn reaction medium 36 is thentransported through an elbow conduit 504 of relatively large diameterand enters the top of deaeration vessel 500. In FIG. 20, elbow conduit504 is shown as connecting orthogonally to the sidewall of vessel shell22 and as comprising a smooth turn through an angle of about 90 degrees.This is merely one convenient configuration; and elbow conduit 504 maybe otherwise in any respect, providing that it usefully connects bubblecolumn reactor 20 with external deaeration vessel 500, as described.Furthermore, elbow conduit 504 may usefully comprise means of flowisolation, such as valves.

In deaeration vessel 500, the gas phase moves upwardly, while the solidand liquid phases move downwardly. The upwardly moving gas phase canre-enter elbow conduit 504 and then escape through opening 502 back intoreaction zone 28. Thus, a counter-current flow of the entering reactionmedium 36 and the exiting disengaged gas can occur at opening 502. Thedeaerated slurry exits deaeration vessel 500 via conduit 506. Deaerationvessel 500 includes a substantially upright, preferably cylindricalsidewall 508 defining a deaeration zone 510. Deaeration zone 510 has aheight “h” and a diameter “d.” It is preferred for elevated opening 502and elbow conduit 504 to have a diameter the same as, or greater than,the diameter “d” of deaeration zone 510. The height “h”, diameter “d”,area, and volume of deaeration zone 510 are preferably substantially thesame as described above with reference to deaeration zone 312 ofdeaeration vessel 300 illustrated in. FIGS. 16-18.

FIGS. 19 and 20 illustrate an embodiment of bubble column reactor 20where the solid product (e.g., crude terephthalic acid) produced inreaction zone 28 is withdrawn from reaction zone 28 via an elevatedoutlet. Withdrawing aerated reaction medium 36 from an elevated locationabove the bottom of bubble column reactor 20 can help avoid accumulationand stagnation of poorly aerated reaction medium 36 at the bottom 52 ofreaction zone 28. According to other aspects of the present invention,the concentrations of oxygen and the oxidizable compound (e.g.,para-xylene) in the reaction medium 36 near the top of reaction medium36 are preferably lower than near the bottom. Thus, withdrawing reactionmedium 36 at an elevated location can increase yield by lowering theamount of unreacted reactants withdrawn from reactor 20. Also, thetemperature of reaction medium 36 varies significantly in the verticaldirection when bubble column reactor 20 is operated with the high STRand the gradients of chemical composition as disclosed herein. Undersuch conditions, the temperature of reaction medium 36 will typicallyhave local minima near the lower end and the upper end of reaction zone28. Near the lower end, the minimum relates to the evaporation ofsolvent near where all or part of the oxidant is admitted. Near theupper end, the minimum is again due to evaporation of solvent, thoughhere due to declining pressure within the reaction medium. In addition,other local minima may occur in between the upper and lower endswherever additional feed or oxidant is admitted to the reaction medium.Thus, there exist one or more temperature maxima, driven by theexothermic heat of oxidation reactions, between the lower end and upperend of reaction zone 28. Withdrawing reaction medium 36 at an elevatedlocation of higher temperature can be particularly advantageous whendownstream processing occurs at higher temperatures, because energycosts associated with heating the withdrawn medium for downstreamprocessing are reduced.

Thus, in a preferred embodiment of the present invention and especiallywhen downstream processing occurs at higher temperatures, reactionmedium 36 is withdrawn from bubble column reactor 20 via an elevatedoutlet(s) positioned above the location(s) where at least 50 weightpercent of the liquid-phase feed stream and/or the gas-phase oxidantstream enter reaction zone 28. More preferably, reaction medium 36 iswithdrawn from bubble column reactor 20 via an elevated outlet(s)positioned above the location(s) where substantially all of theliquid-phase feed stream and/or the gas-phase oxidant stream enterreaction zone 28. Preferably, at least 50 weight percent of thesolid-phase and liquid-phase components withdrawn from bubble columnreactor 20 are withdrawn via an elevated outlet(s). More preferably,substantially all of the solid-phase and liquid-phase componentswithdrawn from bubble column reactor 20 are withdrawn via an elevatedoutlet(s). Preferably, the elevated outlet(s) is located at least aboutID above lower end 52 of reaction zone 28. More preferably, the elevatedoutlet(s) is located at least about 2 D above lower end 52 of reactionzone 28. Most preferably, the elevated outlet(s) is located at least 3 Dabove lower end 52 of reaction zone 28. Given a height “H” of reactionmedium 36, it is preferred for the elevated outlet(s) to be verticallylocated between about 0.2 H and about 0.8 H, more preferably betweenabout 0.3 H and about 0.7 H, and most preferably between 0.4 H and 0.6H. Furthermore, it is preferred that the temperature of reaction medium36 at an elevated outlet from reaction zone 28 is at least 1° C. greaterthan the temperature of reaction medium 36 at lower end 52 of reactionzone 28. More preferably, the temperature of reaction medium 36 at theelevated outlet of reaction zone 28 is in the range of from about 1.5 toabout 16° C. hotter than the temperature of reaction medium 36 at lowerend 52 of reaction zone 28. Most preferably, the temperature of reactionmedium 36 at the elevated outlet of reaction zone 28 is in the range offrom 2 to 12° C. hotter than the temperature of reaction medium 36 atlower end 52 of reaction zone 28.

Referring now to FIG. 21, bubble column reactor 20 is illustrated asincluding an alternative hybrid deaeration vessel 600 positioned at thebottom of reactor 20. In this configuration, aerated reaction medium 36is withdrawn from reaction zone 28 through a relatively large opening602 in the lower end 52 of vessel shell 22. Opening 602 defines the openupper end of deaeration vessel 600. In deaeration vessel 600, the gasphase moves upwardly, while the solid and liquid phases move downwardly.The upwardly moving gas phase can re-enter reaction zone 28 throughopening 602. Thus, a counter-current flow of the entering reactionmedium 36 and the exiting disengaged gas can occur at opening 602. Thedeaerated slurry exits deaeration vessel 600 via conduit 604. Deaerationvessel 600 includes a substantially upright, preferably cylindricalsidewall 606 defining a deaeration zone 608. Deaeration zone 608 has aheight “h” and a diameter “d.” It is preferred for opening 602 to have adiameter the same as, or greater than, the diameter “d” of deaerationzone 608. The height “h”, diameter “d”, area, and volume of deaerationzone 608 are preferably substantially the same as described above withreference to deaeration zone 312 of deaeration vessel 300 illustrated inFIGS. 16-18.

Referring now to FIG. 22, bubble column reactor 20 of FIG. 21 isillustrated as including an alternative oxidant sparger 620. Oxidantsparger 620 includes a ring member 622 and a pair of inlet conduits624,626. Ring member 622 preferably has substantially the sameconfiguration as ring member 202, described above with reference toFIGS. 12-15. Inlet conduits 624,626 extend upwardly through openings inlower head 48 of vessel shell 22 and provide the oxidant stream to ringmember 622.

Referring now to FIG. 23, bubble column reactor 20 of FIG. 21 isillustrated as including a spargerless means for introducing the oxidantstream into reaction zone 28. In the configuration of FIG. 23, theoxidant stream is provided to reactor 20 via oxidant conduits 630,632.Oxidant conduits 630,632 are coupled to respective oxidant openings634,636 in lower head 48 of vessel shell 22. The oxidant stream isintroduced directly into reaction zone 28 via oxidant openings 634,636.Optional impingement plates 638,640 can be provided to deflect the flowof the oxidant stream once it has initially entered reaction zone 28.

As mentioned above, it is preferred for the oxidation reactor to beconfigured and operated in a manner that avoids zones of highconcentration of oxidizable compound in the reaction medium because suchzones can lead to the formation of impurities. One way to improveinitial dispersion of the oxidizable compound (e.g., para-xylene) in thereaction medium is by diluting the oxidizable compound with a liquid.The liquid used to dilute the oxidizable compound can originate from aportion of the reaction medium located a substantial distance from thelocation(s) where the oxidizable compound is fed to the reaction zone.This liquid from a distant portion of the reaction medium can becirculated to a location proximate the location of entry of theoxidizable compound via a flow conduit that is disposed internallyand/or externally to the main reaction vessel.

FIGS. 24 and 25 illustrate two preferred methods of circulating liquidfrom a distant portion of the reaction medium to a location near theinlet of the oxidizable compound using an internal (FIG. 24) or external(FIG. 25) conduit. Preferably, the length of the flow conduit from itsinlet (i.e., opening(s) where the liquid enters the conduit) to itsoutlet (i.e., opening(s) where the liquid is discharge from the conduit)is greater than about 1 meter, more preferably greater than about 3meters, still more preferably greater than about 6 meters, and mostpreferably greater than 9 meters. However, the actual length of theconduit becomes less relevant if the liquid is obtained from a separatevessel, perhaps located immediately above or beside the vessel intowhich the oxidizable compound feed is initially released. Liquid fromany separate vessel containing at least some of the reaction medium is apreferred source for initial dilution of the oxidizable compound.

It is preferred that the liquid flowing through the conduit, whateverthe source, has a lower standing concentration of oxidizable compoundthan the reaction medium immediately adjacent to at least one outlet ofthe conduit. Furthermore, it is preferred that the liquid flowingthrough the conduit has a concentration of oxidizable compound in theliquid phase below about 100,000 ppmw, more preferably below about10,000 ppmw, still more preferably below about 1,000 ppmw and mostpreferably below 100 ppmw, where the concentrations are measured beforeaddition to the conduit of the increment of oxidizable compound feed andof any optional, separate solvent feed. When measured after adding theincrement of oxidizable compound feed and optional solvent feed, it ispreferable that the combined liquid stream entering the reaction mediumhas a concentration of oxidizable compound in the liquid phase belowabout 300,000 ppmw, more preferably below about 50,000 ppmw, and mostpreferably below 10,000 ppmw.

It is desirable to maintain the flow through the conduit at a low enoughrate so that the circulated liquid does suppress the desirable overallgradient of oxidizable compound within the reaction medium. In thisregard, it is preferable that the ratio of the mass of the liquid phasein the reaction zone to which the increment of oxidizable compound isinitially released to the mass flow rate of liquid flowing through theconduit be greater than about 0.3 minutes, more preferably greater thanabout 1 minute, still more preferably between about 2 minutes and about120 minutes, and most preferably between 3 minutes and 60 minutes.

There are many means for compelling the liquid to flow through theconduit. Preferred means include gravity, eductors of all typesemploying either gas or liquid or both as the motive fluid, andmechanical pumps of all types. When using an eductor, one embodiment ofthe invention uses as a motive fluid at least one fluid selected fromthe group consisting of: feed of oxidizable compound (liquid or gas),feed of oxidant (gas), feed of solvent (liquid), and a pumped source ofreaction medium (slurry). Another embodiment uses as a motive fluid atleast two fluids selected from the group consisting of: feed ofoxidizable compound, feed of oxidant, and feed of solvent. Still anotherembodiment uses as a motive fluid a combination feed of oxidizablecompound, feed of oxidant, and feed of solvent.

The appropriate diameter or diameters of the circulation conduit mayvary according to the amount and properties of material being conveyed,the energy available for compelling the flow movement, and considerationof capital cost. It is preferable that the minimum diameter for suchconduit is greater than about 0.02 meters, more preferably between about0.06 meters and about 2 meters, and most preferably between 0.12 and 0.8meters

As noted above, it is desirable to control flow through the conduit incertain preferred ranges. There are many means known in the art toaffect this control by setting an appropriate fixed geometry duringconstruction of the flow conduit. Another preferred embodiment is to usegeometries that are variable during operation, notably including valvesof all sorts and descriptions, including both manual operation andpowered operation by any means, including feed back control loops from asensing element or without. Another preferred means of controlling theflow of the dilution liquid is to vary the energy input between inletand outlet of the conduit. Preferred means include changing the flowrate of one or more motive fluids to an eductor, changing the energyinput to a pump driver, and changing the density difference or elevationdifference when using gravitational force. These preferred means may beused in all combinations as well.

The conduit used for circulation of liquid from the reaction medium maybe of any type known in the art. One embodiment employs a conduitconstructed in whole or part using conventional piping materials.Another embodiment employs a conduit constructed in whole or part usingthe reaction vessel wall as one part of the conduit. A conduit may beconstructed entirely enclosed within the boundaries of the reactionvessel (FIG. 24), or it may be constructed entirely outside the reactionvessel (FIG. 25), or it may comprise sections both within and withoutthe reaction vessel.

The inventors contemplate that, particularly in larger reactors, it maybe desirable to have multiple conduits and of various designs formovement of the liquid through the conduit. Further, it may be desirableto provide multiple outlets at multiple positions on one or all of theconduits. The particulars of the design will balance the desirableoverall gradient in standing concentrations of oxidizable compound withthe desirable initial dilution of oxidizable compound feed, according toother aspects of the current invention.

FIGS. 24 and 25 both illustrate designs that employ a deaeration vesselcoupled to the conduit. This deaeration vessel ensures that the portionof the reaction medium used to dilute the incoming oxidizable compoundis substantially de-aerated slurry. It is now noted, however, that theliquid or slurry used to dilute the incoming oxidizable compound may bein an aerated form as well as a de-aerated form.

The use of a liquid flowing through a conduit to provide dilution of theoxidizable compound feed is particularly useful in bubble columnreactors. Furthermore, in bubble column reactors, a good benefit for theinitial dilution of the oxidizable compound feed can be achieved evenwithout adding the oxidizable compound feed directly into the conduit,providing that the outlet of the conduit lies sufficiently close to theposition of addition of the oxidizable compound. In such an embodiment,it is preferable that the outlet of the conduit be located within about27 conduit outlet diameters of the nearest addition location for theoxidizable compound, more preferably within about 9 conduit outletdiameters, still more preferably within about 3 conduit outletdiameters, and most preferably within 1 conduit outlet diameter.

It has also been discovered that flow eductors can be useful for initialdilution of oxidizable compound feed in oxidation bubble columnsaccording to on embodiment of the present invention, even without theuse of conduits for obtaining dilution liquid from a distant portion ofthe reaction medium. In such cases, the eductor is located within thereaction medium and has an open pathway from the reaction medium intothe throat of the eductor, where low pressure draws in adjacent reactionmedium. Examples of two possible eductor configurations are illustratedin FIGS. 26 and 27. In a preferred embodiment of these eductors, thenearest location of feeding oxidizable compound is within about 4meters, more preferably within about 1 meter and most preferably 0.3meters of the throat of the eductor. In another embodiment, theoxidizable compound is fed under pressure as a motive fluid. In stillanother embodiment, either the solvent or the oxidant is fed underpressure as additional motive fluid along with the oxidizable compound.In yet another embodiment, both the solvent and ant oxidant are fedunder pressure as additional motive fluid along with the oxidizablecompound.

The inventors contemplate that, particularly in larger reactors, it maybe desirable to have multiple eductors and of various designs situatedat various positions within the reaction medium. The particulars of thedesign will balance the desirable overall gradient in standingconcentrations of the oxidizable compound with the desirable initialdilution of the oxidizable compound feed, according to other aspects ofthe current invention. In addition, the inventors contemplate that theoutlet flow plumes from an eductor may be oriented in any direction.When multiple eductors are used, each eductor may be orientedseparately, again in any direction.

As mentioned above, certain physical and operational features of bubblecolumn reactor 20, described above with reference to FIGS. 1-27, providefor vertical gradients in the pressure, temperature, and reactant (i.e.,oxygen and oxidizable compound) concentrations of reaction medium 36. Asdiscussed above, these vertical gradients can provide for a moreeffective and economical oxidation process as compared to conventionaloxidations processes, which favor a well-mixed reaction medium ofrelatively uniform pressure, temperature, and reactant concentrationthroughout. The vertical gradients for oxygen, oxidizable compound(e.g., para-xylene), and temperature made possible by employing anoxidation system in accordance with an embodiment of the presentinvention will now be discussed in greater detail.

Referring now to FIG. 28, in order to quantify the reactantconcentration gradients existing in reaction medium 36 during oxidationin bubble column reactor 20, the entire volume of reaction medium 36 canbe theoretically partitioned into 30 discrete horizontal slices of equalvolume. FIG. 28 illustrates the concept of dividing reaction medium 36into 30 discrete horizontal slices of equal volume. With the exceptionof the highest and lowest horizontal slices, each horizontal slice is adiscrete volume bounded on its top and bottom by imaginary horizontalplanes and bounded on its sides by the wall of reactor 20. The highesthorizontal slice is bounded on its bottom by an imaginary horizontalplane and on its top by the upper surface of reaction medium 36. Thelowest horizontal slice is bounded on its top by an imaginary horizontalplane and on its bottom by the bottom of the vessel shell. Once reactionmedium 36 has been theoretically partitioned into 30 discrete horizontalslices of equal volume, the time-averaged and volume-averagedconcentration of each horizontal slice can then be determined. Theindividual horizontal slice having the maximum concentration of all 30horizontal slices can be identified as the “C-max horizontal slice.” Theindividual horizontal slice located above the C-max horizontal slice andhaving the minimum concentration of all horizontal slices located abovethe C-max horizontal slice can be identified as the “C-min horizontalslice.” The vertical concentration gradient can then be calculated asthe ratio of the concentration in the C-max horizontal slice to theconcentration in the C-min horizontal slice.

With respect to quantifying the oxygen concentration gradient, whenreaction medium 36 is theoretically partitioned into 30 discretehorizontal slices of equal volume, an O₂-max horizontal slice isidentified as having the maximum oxygen concentration of all the 30horizontal slices and an O₂-min horizontal slice is identified as havingthe minimum oxygen concentration of the horizontal slices located abovethe O₂-max horizontal slice. The oxygen concentrations of the horizontalslices are measured in the gas phase of reaction medium 36 on atime-averaged and volume-averaged molar wet basis. It is preferred forthe ratio of the oxygen concentration of the O₂-max horizontal slice tothe oxygen concentration of the O₂-min horizontal slice to be in therange of from about 2:1 to about 25:1, more preferably in the range offrom about 3:1 to about 15:1, and most preferably in the range of from4:1 to 10:1.

Typically, the O₂-max horizontal slice will be located near the bottomof reaction medium 36, while the O₂-min horizontal slice will be locatednear the top of reaction medium 36. Preferably, the O₂-min horizontalslice is one of the 5 upper-most horizontal slices of the 30 discretehorizontal slices. Most preferably, the O₂-min horizontal slice is theupper-most one of the 30 discrete horizontal slices, as illustrated inFIG. 28. Preferably, the O₂-max horizontal slice is one of the 10lower-most horizontal slices of the 30 discrete horizontal slices. Mostpreferably, the O₂-max horizontal slice is one of the 5 lower-mosthorizontal slices of the 30 discrete horizontal slices. For example,FIG. 28 illustrates the O₂-max horizontal slice as the third horizontalslice from the bottom of reactor 20. It is preferred for the verticalspacing between the O₂-min and O₂-max horizontal slices to be at leastabout 2 W, more preferably at least about 4 W, and most preferably atleast 6 W. It is preferred for the vertical spacing between the O₂-minand O₂-max horizontal slices to be at least about 0.2 H, more preferablyat least about 0.4 H, and most preferably at least 0.6 H Thetime-averaged and volume-averaged oxygen concentration, on a wet basis,of the O₂-min horizontal slice is preferably in the range of from about0.1 to about 3 mole percent, more preferably in the range of from about0.3 to about 2 mole percent, and most preferably in the range of from0.5 to 1.5 mole percent. The time-averaged and volume-averaged oxygenconcentration of the O₂-max horizontal slice is preferably in the rangeof from about 4 to about 20 mole percent, more preferably in the rangeof from about 5 to about 15 mole percent, and most preferably in therange of from 6 to 12 mole percent. The time-averaged concentration ofoxygen, on a dry basis, in the gaseous effluent discharged from reactor20 via gas outlet 40 is preferably in the range of from about 0.5 toabout 9 mole percent, more preferably in the range of from about 1 toabout 7 mole percent, and most preferably in the range of from 1.5 to 5mole percent.

Because the oxygen concentration decays so markedly toward the top ofreaction medium 36, it is desirable that the demand for oxygen bereduced in the top of reaction medium 36. This reduced demand for oxygennear the top of reaction medium 36 can be accomplished by creating avertical gradient in the concentration of the oxidizable compound (e.g.,para-xylene), where the minimum concentration of oxidizable compound islocated near the top of reaction medium 36.

With respect to quantifying the oxidizable compound (e.g., para-xylene)concentration gradient, when reaction medium 36 is theoreticallypartitioned into 30 discrete horizontal slices of equal volume, anOC-max horizontal slice is identified as having the maximum oxidizablecompound concentration of all the 30 horizontal slices and an OC-minhorizontal slice is identified as having the minimum oxidizable compoundconcentration of the horizontal slices located above the OC-maxhorizontal slice. The oxidizable compound concentrations of thehorizontal slices are measured in the liquid phase on a time-averagedand volume-averaged mass fraction basis. It is preferred for the ratioof the oxidizable compound concentration of the OC-max horizontal sliceto the oxidizable compound concentration of the OC-min horizontal sliceto be greater than about 5:1, more preferably greater than about 10:1,still more preferably greater than about 20:1, and most preferably inthe range of from 40:1 to 1000:1.

Typically, the OC-max horizontal slice will be located near the bottomof reaction medium 36, while the OC-min horizontal slice will be locatednear the top of reaction medium 36. Preferably, the OC-min horizontalslice is one of the 5 upper-most horizontal slices of the 30 discretehorizontal slices. Most preferably, the OC-min horizontal slice is theupper-most one of the 30 discrete horizontal slices, as illustrated inFIG. 28. Preferably, the OC-max horizontal slice is one of the 10lower-most horizontal slices of the 30 discrete horizontal slices. Mostpreferably, the OC-max horizontal slice is one of the 5 lower-mosthorizontal slices of the 30 discrete horizontal slices. For example,FIG. 28 illustrates the OC-max horizontal slice as the fifth horizontalslice from the bottom of reactor 20. It is preferred for the verticalspacing between the OC-min and OC-max horizontal slices to be at leastabout 2 W, where “W” is the maximum width of reaction medium 36. Morepreferably, the vertical spacing between the OC-min and OC-maxhorizontal slices is at least about 4 W, and most preferably at least 6W. Given a height “H” of reaction medium 36, it is preferred for thevertical spacing between the OC-min and OC-max horizontal slices to beat least about 0.2 H, more preferably at least about 0.4 H, and mostpreferably at least 0.6 H.

The time-averaged and volume-averaged oxidizable compound (e.g.,para-xylene) concentration in the liquid phase of the OC-min horizontalslice is preferably less than about 5,000 ppmw, more preferably lessthan about 2,000 ppmw, still more preferably less than about 400 ppmw,and most preferably in the range of from 1 ppmw to 100 ppmw. Thetime-averaged and volume-averaged oxidizable compound concentration inthe liquid phase of the OC-max horizontal slice is preferably in therange of from about 100 ppmw to about 10,000 ppmw, more preferably inthe range of from about 200 ppmw to about 5,000 ppmw, and mostpreferably in the range of from 500 ppmw to 3,000 ppmw.

Although it is preferred for bubble column reactor 20 to providevertical gradients in the concentration of the oxidizable compound, itis also preferred that the volume percent of reaction medium 36 havingan oxidizable compound concentration in the liquid phase above 1,000ppmw be minimized. Preferably, the time-averaged volume percent ofreaction medium 36 having an oxidizable compound concentration in theliquid phase above 1,000 ppmw is less than about 9 percent, morepreferably less than about 6 percent, and most preferably less than 3percent. Preferably, the time-averaged volume percent of reaction medium36 having an oxidizable compound concentration in the liquid phase above2,500 ppmw is less than about 1.5 percent, more preferably less thanabout 1 percent, and most preferably less than 0.5 percent. Preferably,the time-averaged volume percent of reaction medium 36 having anoxidizable compound concentration in the liquid phase above 10,000 ppmwis less than about 0.3 percent, more preferably less than about 0.1percent, and most preferably less than 0.03 percent. Preferably, thetime-averaged volume percent of reaction medium 36 having an oxidizablecompound concentration in the liquid phase above 25,000 ppmw is lessthan about 0.03 percent, more preferably less than about 0.015 percent,and most preferably less than 0.007 percent. The inventors note that thevolume of reaction medium 36 having the elevated levels of oxidizablecompound need not lie in a single contiguous volume. At many times, thechaotic flow patterns in a bubble column reaction vessel producesimultaneously two or more continuous but segregated portions ofreaction medium 36 having the elevated levels of oxidizable compound. Ateach time used in the time averaging, all such continuous but segregatedvolumes larger than 0.0001 volume percent of the total reaction mediumare added together to determine the total volume having the elevatedlevels of oxidizable compound concentration in the liquid phase.

In addition to the concentration gradients of oxygen and oxidizablecompound, discussed above, it is preferred for a temperature gradient toexist in reaction medium 36. Referring again to FIG. 28, thistemperature gradient can be quantified in a manner similar to theconcentration gradients by theoretically partitioning reaction medium 36into 30 discrete horizontal slices of equal volume and measuring thetime-averaged and volume-averaged temperature of each slice. Thehorizontal slice with the lowest temperature out of the lowest 15horizontal slices can then be identified as the T-min horizontal slice,and the horizontal slice located above the T-min horizontal slice andhaving the maximum temperature of all the slices above the T-minhorizontal slice can then be identified as the “T-max horizontal slice.”It is preferred for the temperature of the T-max horizontal slice be atleast about 1° C. higher than the temperature of the T-min horizontalslice. More preferably the temperature of the T-max horizontal slice isin the range of from about 1.25 to about 12° C. higher than thetemperature of the T-min horizontal slice. Most preferably thetemperature of the T-max horizontal slice is in the range of from 2 to8° C. higher than the temperature of the T-min horizontal slice. Thetemperature of the T-max horizontal slice is preferably in the range offrom about 125 to about 200° C., more preferably in the range of fromabout 140 to about 180° C., and most preferably in the range of from 150to 170° C.

Typically, the T-max horizontal slice will be located near the center ofreaction medium 36, while the T-min horizontal slice will be locatednear the bottom of reaction medium 36. Preferably, the T-min horizontalslice is one of the 10 lower-most horizontal slices of the 15 lowesthorizontal slices. Most preferably, the T-min horizontal slice is one ofthe 5 lower-most horizontal slices of the 15 lowest horizontal slices.For example, FIG. 28 illustrates the T-min horizontal slice as thesecond horizontal slice from the bottom of reactor 20. Preferably, theT-max horizontal slice is one of the 20 middle horizontal slices of the30 discrete horizontal slices. Most preferably, the T-min horizontalslice is one of the 14 middle horizontal slices of the 30 discretehorizontal slices. For example, FIG. 28 illustrates the T-max horizontalslice as the twentieth horizontal slice from the bottom of reactor 20(i.e., one of the middle 10 horizontal slices). It is preferred for thevertical spacing between the T-min and T-max horizontal slices to be atleast about 2 W, more preferably at least about 4 W, and most preferablyat least 6 W. It is preferred for the vertical spacing between the T-minand T-max horizontal slices to be at least about 0.2 H, more preferablyat least about 0.4 H, and most preferably at least 0.6 H.

As discussed above, when a vertical temperature gradient exists inreaction medium 36, it can be advantageous to withdraw reaction medium36 at an elevated location where the temperature of reaction medium ishighest, especially when the withdrawn product is subjected to furtherdownstream processing at higher temperatures. Thus, when reaction medium36 is withdrawn from reaction zone 28 via one or more elevated outlets,as illustrated in FIGS. 19 and 20, it is preferred for the elevatedoutlet(s) to be located near the T-max horizontal slice. Preferably, theelevated outlet is located within 10 horizontal slices of the T-maxhorizontal slice, more preferably within 5 horizontal slices of theT-max horizontal slice, and most preferably within 2 horizontal slicesof the T-max horizontal slice.

It is now noted that many of the inventive features described herein canbe employed in multiple oxidation reactor systems—not just systemsemploying a single oxidation reactor. In addition, certain inventivefeatures described herein can be employed in mechanically-agitatedand/or flow-agitated oxidation reactors—not just bubble-agitatedreactors (i.e., bubble column reactors). For example, the inventors havediscovered certain advantages associated with staging/varying oxygenconcentration and/or oxygen consumption rate throughout the reactionmedium. The advantages realized by the staging of oxygenconcentration/consumption in the reaction medium can be realized whetherthe total volume of the reaction medium is contained in a single vesselor in multiple vessels. Further, the advantages realized by the stagingof oxygen concentration/consumption in the reaction medium can berealized whether the reaction vessel(s) is mechanically-agitated,flow-agitated, and/or bubble-agitated.

One way of quantifying the degree of staging of oxygen concentrationand/or consumption rate in a reaction medium is to compare two or moredistinct 20-percent continuous volumes of the reaction medium. These20-percent continuous volumes need not be defined by any particularshape. However, each 20-percent continuous volume must be formed of acontiguous volume of the reaction medium (i.e., each volume is“continuous”), and the 20-percent continuous volumes must not overlapone another (i.e., the volumes are “distinct”). FIGS. 29-31 illustratethat these distinct 20-percent continuous volumes can be located in thesame reactor (FIG. 29) or in multiple reactors (FIGS. 30 and 31). It isnoted that the reactors illustrated in FIGS. 29-31 can bemechanically-agitated, flow-agitated, and/or bubble-agitated reactors.In one embodiment, it is preferred for the reactors illustrated in FIGS.29-31 to be bubble-agitated reactors (i.e., bubble column reactors).

Referring now to FIG. 29, reactor 20 is illustrated as containing areaction medium 36. Reaction medium 36 includes a first distinct20-percent continuous volume 37 and a second distinct 20-percentcontinuous volume 39.

Referring now to FIG. 30, a multiple reactor system is illustrated asincluding a first reactor 720 a and a second reactor 720 b. Reactors 720a,b cooperatively contain a total volume of a reaction medium 736. Firstreactor 720 a contains a first reaction medium portion 736 a, whilesecond reactor 720 b contains a second reaction medium portion 736 b. Afirst distinct 20-percent continuous volume 737 of reaction medium 736is shown as being defined within first reactor 720 a, while a seconddistinct 20-percent continuous volume 739 of reaction medium 736 isshown as being defined within second reactor 720 b.

Referring now to FIG. 31, a multiple reactor system is illustrated asincluding a first reactor 820 a, a second reactor 820 b, and a thirdreactor 820 c. Reactors 820 a,b,c cooperatively contain a total volumeof a reaction medium 836. First reactor 820 a contains a first reactionmedium portion 836 a; second reactor 820 b contains a second reactionmedium portion 836 b; and third reactor 820 c contains a third reactionmedium portion 836 c. A first distinct 20-percent continuous volume 837of reaction medium 836 is shown as being defined within first reactor820 a; a second distinct 20-percent continuous volume 839 of reactionmedium 836 is shown as being defined within second reactor 820 b; and athird distinct 20-percent continuous volume 841 of reaction medium 836is show as being defined within third reactor 820 c.

The staging of oxygen availability in the reaction medium can bequantified by referring to the 20-percent continuous volume of reactionmedium having the most abundant mole fraction of oxygen in the gas phaseand by referring to the 20-percent continuous volume of reaction mediumhaving the most depleted mole fraction of oxygen in the gas phase. Inthe gas phase of the distinct 20-percent continuous volume of thereaction medium containing the highest concentration of oxygen in thegas phase, the time-averaged and volume-averaged oxygen concentration,on a wet basis, is preferably in the range of from about 3 to about 18mole percent, more preferably in the range of from about 3.5 to about 14mole percent, and most preferably in the range of from 4 to 10 molepercent. In the gas phase of the distinct 20-percent continuous volumeof the reaction medium containing the lowest concentration of oxygen inthe gas phase, the time-averaged and volume-averaged oxygenconcentration, on a wet basis, is preferably in the range of from about0.3 to about 5 mole percent, more preferably in the range of from about0.6 to about 4 mole percent, and most preferably in the range of from0.9 to 3 mole percent. Furthermore, the ratio of the time-averaged andvolume-averaged oxygen concentration, on a wet basis, in the mostabundant 20-percent continuous volume of reaction medium compared to themost depleted 20-percent continuous volume of reaction medium ispreferably in the range of from about 1.5:1 to about 20:1, morepreferably in the range of from about 2:1 to about 12:1, and mostpreferably in the range of from 3:1 to 9:1.

The staging of oxygen consumption rate in the reaction medium can bequantified in terms of an oxygen-STR, initially described above.Oxygen-STR was previously describe in a global sense (i.e., from theperspective of the average oxygen-STR of the entire reaction medium);however, oxygen-STR may also be considered in a local sense (i.e., aportion of the reaction medium) in order to quantify staging of theoxygen consumption rate throughout the reaction medium.

The inventors have discovered that it is very useful to cause theoxygen-STR to vary throughout the reaction medium in general harmonywith the desirable gradients disclosed herein relating to pressure inthe reaction medium and to the mole fraction of molecular oxygen in thegas phase of the reaction medium. Thus, it is preferable that the ratioof the oxygen-STR of a first distinct 20-percent continuous volume ofthe reaction medium compared to the oxygen-STR of a second distinct20-percent continuous volume of the reaction medium be in the range offrom about 1.5:1 to about 20:1, more preferably in the range of fromabout 2:1 to about 12:1, and most preferably in the range of from 3:1 to9:1. In one embodiment the “first distinct 20-percent continuous volume”is located closer than the “second distinct 20-percent continuousvolume” to the location where molecular oxygen is initially introducedinto the reaction medium. These large gradients in oxygen-STR aredesirable whether the partial oxidation reaction medium is contained ina bubble column oxidation reactor or in any other type of reactionvessel in which gradients are created in pressure and/or mole fractionof molecular oxygen in the gas phase of the reaction medium (e.g., in amechanically agitated vessel having multiple, vertically disposedstirring zones achieved by using multiple impellers having strong radialflow, possibly augmented by generally horizontal baffle assemblies, withoxidant flow rising generally upwards from a feed near the lower portionof the reaction vessel, notwithstanding that considerable back-mixing ofoxidant flow may occur within each vertically disposed stirring zone andthat some back-mixing of oxidant flow may occur between adjacentvertically disposed stirring zones). That is, when a gradient exists inthe pressure and/or mole fraction of molecular oxygen in the gas phaseof the reaction medium, the inventors have discovered that it isdesirable to create a similar gradient in the chemical demand fordissolved oxygen by the means disclosed herein.

A preferred means of causing the local oxygen-STR to vary is bycontrolling the locations of feeding the oxidizable compound and bycontrolling the mixing of the liquid phase of the reaction medium tocontrol gradients in concentration of oxidizable compound according toother disclosures of the present invention. Other useful means ofcausing the local oxygen-STR to vary include causing variation inreaction activity by causing local temperature variation and by changingthe local mixture of catalyst and solvent components (e.g., byintroducing an additional gas to cause evaporative cooling in aparticular portion of the reaction medium and by adding a solvent streamcontaining a higher amount of water to decrease activity in a particularportion of the reaction medium).

As discussed above with reference to FIGS. 30 and 31, the partialoxidation reaction can be usefully conducted in multiple reactionvessels wherein at least a portion, preferably at least 25 percent, morepreferably at least 50 percent, and most preferable at least 75 percent,of the molecular oxygen exiting from a first reaction vessel isconducted to one or more subsequent reaction vessels for consumption ofan additional increment, preferably more than 10 percent, morepreferably more than 20 percent, and most preferably more than 40percent, of the molecular oxygen exiting the first/upstream reactionvessel. When using such a series flow of molecular oxygen from onereactor to others, it is desirable that the first reaction vessel isoperated with a higher reaction intensity than at least one of thesubsequent reaction vessels, preferably with the ratio of thevessel-average-oxygen-STR within the first reaction vessel to thevessel-average-oxygen-STR within the subsequent reaction vessel in therange of from about 1.5:1 to about 20:1, more preferably in the range offrom about 2:1 to about 12:1, and most preferably in the range of from3:1 to 9:1.

As discussed above, all types of first reaction vessel (e.g.; bubblecolumn, mechanically-agitated, back-mixed, internally staged, plug flow,and so on) and all types of subsequent reaction vessels, which may ornot be of different type than the first reaction vessel, are useful forseries flow of molecular oxygen to subsequent reaction vessels withaccording to the present invention. The means of causing thevessel-average-oxygen-STR to decline within subsequent reaction vesselsusefully include reduced temperature, reduced concentrations ofoxidizable compound, and reduced reaction activity of the particularmixture of catalytic components and solvent (e.g., reduced concentrationof cobalt, increased concentration of water, and addition of a catalyticretardant such as small quantities of ionic copper).

In flowing from the first reaction vessel to a subsequent reactionvessel, the oxidant stream may be treated by any means known in the artsuch as compression or pressure reduction, cooling or heating, andremoving mass or adding mass of any amount or any type. However, the useof declining vessel-average-oxygen-STR in subsequent reaction vessels isparticularly useful when the absolute pressure in the upper portion ofthe first reaction vessel is less than about 2.0 megapascal, morepreferably less than about 1.6 megapascal, and most preferably less than1.2 megapascal. Furthermore, the use of decliningvessel-average-oxygen-STR in subsequent reaction vessels is particularlyuseful when the ratio of the absolute pressure in the upper portion ofthe first reaction vessel compared to the absolute pressure in the upperportion of at least one subsequent reaction vessel is in the range fromabout 0.5:1 to 6:1, more preferably in a range from about 0.6:1 to about4:1, and most preferably in a range from 0.7:1 to 2:1. Pressurereductions in subsequent vessels below these lower bounds overly reducethe availability of molecular oxygen, and pressure increases above theseupper bounds are increasingly costly compared to using a fresh supply ofoxidant.

When using series flow of molecular oxygen to subsequent reactionvessels having declining vessel-average-oxygen-STR, fresh feed streamsof oxidizable compound, solvent and oxidant may flow into subsequentreaction vessels and/or into the first reaction vessel. Flows of theliquid phase and the solid phase, if present, of the reaction medium mayflow in any direction between reaction vessels. All or part of the gasphase leaving the first reaction vessel and entering a subsequentreaction vessel may flow separated from or commingled with portions ofthe liquid phase or the solid phase, if present, of the reaction mediumfrom the first reaction vessel. A flow of product stream comprisingliquid phase and solid phase, if present, may be withdrawn from thereaction medium in any reaction vessel in the system.

Referring again to FIGS. 1-29, in one embodiment of the presentinvention, oxidation bubble column reactor 20 has a significantly higherproduction rate than conventional oxidation bubble column reactors,particularly conventional bubble column reactors used to produceterephthalic acid. In order to provide increased production rates, thesize of bubble column reactor 20 must be increased. However, thenaturally-convected, multi-phase fluid flow dynamics of the reactionmedium in such a scaled-up bubble column reactor are significantly,troublesomely different than the flow dynamics in smaller conventionalreactors. It has been discovered that certain design and operatingparameters are important for the proper functionality of high productionrate scaled-up oxidation bubble column reactors.

When bubble column 20 is a high production rate scaled-up oxidationbubble column reactor in accordance with one embodiment of the presentinvention, the height “H” of reaction medium 36 is preferably at leastabout 30 meters, more preferable in the range of from about 35 to about70 meters, and most preferably in the range of from 40 to 60 meters. Thedensity and height of reaction medium 36 in scaled-up bubble columnreactor 20 cause a significant pressure differential between the top ofreaction medium 36 and the bottom of reaction medium 36. Preferably thepressure differential between the top and bottom of reaction medium 36is at least about 1 bar, more preferably at least about 1.4 bar, andmost preferably in the range of from 1.6 to 3 bar. The maximum width “W”of reaction medium 36 is preferably at least about 2.5 meters, morepreferable in the range of from about 3 to about 20 meters, still morepreferably in the range of from about 3.25 to about 12 meters, and mostpreferably in the range of from 4 to 10 meters. The H:W ratio ofreaction medium 36 is preferably at least about 6:1, more preferably inthe range of from about 8:1 to about 20:1, and most preferably in therange of from 9:1 to 15:1. The total volume of reaction medium 36 and/orreaction zone 28 is preferably at least about 250 cubic meters, morepreferably at least about 500 cubic meters, and most preferably at least1,000 cubic meters.

During operation of scaled-up oxidation bubble column reactor 20, it ispreferred for the time-averaged superficial velocity of the gas phase ofreaction medium 36 at one-quarter height, half height, and/orthree-quarter height to be in the range of from about 0.8 to about 5meters per second, more preferably in the range of from about 0.9 toabout 3 meters per second, and most preferably in the range of from 1 to2 meters per second. When scaled-up oxidation bubble column reactor 20is employed to produce terephthalic acid via partial oxidation ofpara-xylene, it is preferred for para-xylene to be fed to reaction zone28 at a rate of at least about 11,000 kilograms per hour, morepreferably at a rate in the range of from about 20,000 to about 100,000kilograms per hour, and most preferably in the range of from 30,000 to80,000 kilograms per hour. The terephthalic acid production rate ofscaled-up bubble column 20 is preferably at least about 400 tons perday, more preferably at least about 700 tons per day, and mostpreferably in the range of from 750 to 3,000 tons per day. Other designand operating parameters of scaled-up oxidation bubble column reactor 20can be substantially the same as initially described above withreference to FIGS. 1-29.

It has been discovered that varying the horizontal cross-sectional areaof the reaction zone of a bubble column reactor along the height of thereactor can help improve the fluid flow dynamics of the reaction medium,especially in the scaled-up oxidation bubble column reactor designsdiscussed above. Referring now to FIG. 32, in one embodiment of thepresent invention, vessel shell 22 of oxidation bubble column reactor 20includes a broad lower section 23, a narrow upper section 25, and atransition section 27. Preferably, lower and upper sections 23,25 aresubstantially cylindrical in shape and are aligned along a commoncentral upright axis. Transition section 27 may have many suitableshapes (e.g., a horizontal planar shape, a 2:1 elliptical shape, ahemispherical shape, and so on). Preferably, transition section 27 is agenerally frustoconical member that transitions vessel shell 22 frombroad lower section 23 to narrow upper section 25. Lower section 23 ofvessel shell 22 defines a broad lower reaction zone 29. Upper section 25of vessel shell 22 defines a narrow upper reaction zone 31. Transitionsection 27 defines a transition zone located between lower and upperreaction zones 29,31. Lower reaction zone 29, upper reaction zone 31,and the transition zone cooperatively form the total reaction zone ofoxidation bubble column reactor 20 that receives multi-phase reactionmedium 36.

In a preferred embodiment of the present invention, the maximumhorizontal cross-sectional area of reaction medium 36 in lower reactionzone 29 is at least about 10 percent greater than the minimumcross-sectional area of reaction medium 36 in upper reaction zone 31.More preferably, the maximum horizontal cross-sectional area of reactionmedium 36 in lower reaction zone 29 is in the range of from about 25 toabout 210 percent greater than the minimum horizontal cross-sectionalarea of reaction medium 36 in upper reaction zone 31. Most preferably,the maximum horizontal cross-sectional area of reaction medium 36 inlower reaction zone 29 is in the range of from 35 to 160 percent greaterthan the minimum horizontal cross-sectional area of reaction medium 36in upper reaction zone 31.

As illustrated in FIG. 32, lower reaction zone 29 has a maximum diameter“D_(l)” that is greater than the minimum diameter “D_(u)” of upperreaction zone 31. Preferably, D_(l) is at least about 5 percent greaterthan D_(u). More preferably, D_(l) is in the range of from about 10 toabout 100 percent greater than about D_(u). Most preferably, D_(l) is inthe range of from 15 to 50 percent greater than D_(u). FIG. 32 alsoillustrates that lower reaction zone 29 has a maximum height “L_(l)”,upper reaction zone 31 has a maximum height “L_(u)”, and the transitionzone has a maximum height “L_(t)”. It should be noted that the height ofthe portion of reaction medium 36 contained in lower reaction zone 29 isL_(l) and the height of the portion of reaction medium 36 contained inthe transition zone is L_(t). In one embodiment, the height of theportion of reaction medium 36 located in upper reaction zone 31 isL_(u). However, in certain instances the height of reaction medium 36 inupper reaction zone 31 may be less than L_(u). In other instances, thetotal height of reaction medium 36 may extend above the top end 50 ofupper reaction zone 31 (i.e. the total height of reaction medium 36 ismore than the sum of L_(l) plus L_(t) plus L_(u). Preferably, the totalheight of the reaction medium 36 is within 50, 25, or 10 percent ofL_(u) measured either up or down from the top end 50 of upper reactionzone 31. Preferably, oxidation bubble column reactor 20 has anL_(l):L_(u) ratio in the range of from about 0.05:1 to about 5:1, morepreferably in the range of from about 0.1:1 to about 2.5:1, and mostpreferably in the range of from 0.15:1 to 1.5:1. Preferably, oxidationbubble column reactor 20 has an L_(l):D_(l) ratio greater than about0.5:1, more preferably in the range of from about 1:1 to about 10:1, andmost preferably in the range of from 1.5:1 to 8:1. Oxidation bubblecolumn reactor 20 preferably has an L_(u):D_(u) ratio greater than about2:1, more preferably in the range of from about 2.5:1 to about 20:1, andmost preferably in the range of from 3:1 to 15:1. In a preferredembodiment of the present invention L_(l) is at least about 2 meters,more preferably L_(l) is in the range of from about 4 to about 50meters, and most preferably in the range of from 6 to 40 meters.Preferably, L_(u) is at least about 4 meters, more preferably in therange of from about 5 to about 80 meters, and most preferably in therange of from 10 to 60 meters. Preferably, D_(l) is in the range of fromabout 2 to about 12 meters, more preferably in the range of from about3.1 to about 10 meters, and most preferably in the range of from 4 to 9meters.

FIG. 32 also illustrates that oxidation bubble column reactor 20 has adisengagement section 26 located above upper reaction section 31.Disengagement section 26 defines disengagement zone 30. As shown in FIG.32, disengagement zone 30 has a maximum height “Y” and a maximum width“X”. It is preferred for oxidation bubble column reactor 20 to have anX:D_(l) ratio in the range of from about 0.8:1 to about 4:1, mostpreferably in the range of from 1.1:1 to 2:1. Oxidation bubble columnreactor 20 preferably has an L_(u):Y ratio in the range of from about1:1 to about 10:1, most preferably in the range from 2:1 to 5:1.Oxidation bubble column reactor 20 preferably has an L_(l);Y ratio inthe range of from about 0.4:1 to about 8:1, most preferably in the rangefrom 0.6:1 to 4:1. Transition section 27 of vessel shell 22 has amaximum diameter “D_(l)” adjacent lower section 23 and minimum diameter“D_(u)” adjacent upper section 25. Oxidation bubble column reactor 20preferably has an L_(t):D_(l) ratio in the range of from about 0.05:1 toabout 5:1, most preferably in the range of from 0.1:1 to 2:1.

The vertically-varying horizontal cross-sectional area of oxidationbubble column reactor 20 illustrated in FIG. 32 provides the portion ofreaction medium 36 contained in upper reaction zone 31 with a highersuperficial gas velocity and higher gas hold-up than the portion ofreaction medium 36 contained in lower reaction zone 29. Preferably, thetime-averaged superficial gas velocity of the portion of reaction medium36 contained in upper reaction zone 31 at half height of the reactionmedium in upper reaction zone 31 is at least about 15 percent greaterthan the time-averaged superficial gas velocity of the portion ofreaction medium 36 contained in lower reaction zone 29 at half height ofthe reaction medium in lower reaction zone 29. More preferably, theportion of reaction medium 36 contained in upper reaction zone 31 has atime-averaged superficial gas velocity at half-height of the reactionmedium in upper reaction zone 31 that is in the range of from about 25to about 210 percent greater than the time-averaged superficial gasvelocity of the portion of reaction medium 36 contained in lowerreaction zone 29 at half height of the reaction medium in lower reactionzone 29. Most preferably, the time-averaged superficial gas velocity ofthe portion of reaction medium 36 contained in upper reaction zone 31 athalf height of the reaction medium in upper reaction zone 31 is in therange of from 35 to 160 percent greater than the time-averagedsuperficial gas velocity of the portion of reaction medium 36 containedin lower reaction zone 29 at half height of the reaction medium in lowerreaction zone 29. Preferably, the time-averaged and volume-averaged gashold-up of the portion of reaction medium 36 contained in upper reactionzone 31 is at least about 4 percent greater than the time-averaged andvolume-average gas hold-up of the portion of reaction medium 36contained in a lower reaction zone 29. More preferably, thetime-averaged and volume-averaged gas hold-up of the portion of reactionmedium 36 contained in upper reaction zone 31 is in the range of fromabout 6 to about 80 percent greater than the time-averaged andvolume-averaged gas hold-up of the portion of reaction medium 36contained in lower reaction zone 29. Most preferably, the time-averagedand volume-averaged gas hold-up of the portion of reaction medium 36contained in upper reaction zone 31 is in the range of from 8 to 70percent greater than the time-averaged and volume-averaged gas hold-upof the portion of reaction medium 36 contained in lower reaction zone29.

Although FIG. 32 illustrates a very specific two-stage, cylindricalsidewall bubble column design, it should be understood that many otherdesigns may fall within the ambit of this embodiment of the invention.For example, the narrow upper and broad lower sections of the reactormay be formed of one or more sloped sidewalls and/or a plurality ofstepped-diameter sidewall segments. In any case, the claims, not thedrawing, capture the essence of this embodiment.

As mentioned above, in certain instances, it may be desirable to employlarger bubble column reactors in order to permit higher production ratesfor a single reactor. However, as the size of bubble column reactorsincreases, the fluid flow behavior of the multi-phase reaction mediumcontained therein changes significantly from the flow behavior ofsmaller reactors. It has been discovered that the changing fluid flowbehavior of larger bubble column reactors can be counteracted bycontacting the multi-phase reaction medium contained in the bubblecolumn reactor with additional upright surfaces. Accordingly, FIGS.33-44 illustrate various ways of providing additional upright surfacearea in the reaction zone of a scaled-up bubble column reactor. Each ofthe bubble column reactors illustrated in FIGS. 33-44 include one ormore upright internal members contained in the reaction zone of thebubble column reactor. These upright internal members are provided inaddition to the upright pressure-containing sidewalls of the reactor. Aspreviously discussed, the reaction medium contained in the bubble columnreactor has a maximum height “H” and a maximum width “W”. The minimumwidth of the reaction medium occurring above a height of H/5 is referredto herein as “Wmin”. In a preferred embodiment of the present inventionthe total upright surface area of the bubble column reactor thatcontacts the reaction medium is preferably greater than about 3.25WminH, more preferably in the range of from about 3.5 WminH to about 20WminH, still more preferably in the range of from about 4 WminH to about15 WminH, and most preferably in the range of from 5 WminH to 10 WminH.This total upright surface area that contacts the reaction mediumincludes all area of upright surfaces, including upright surfaces of thepressure-containing reactor sidewall and upright surfaces of anyinternal members present within the reaction zone. As used herein theterm “upright” denotes less than 45° from vertical. Preferably, theupright surfaces contacting the reaction medium in the bubble columnreactor extend at an angle within 30° of vertical, more preferablywithin 15° of vertical, still more preferable within 5° of vertical, andmost preferably substantially vertical.

It is further preferred that the total amount of upright surface areacontacting the reaction medium that is attributable to thenon-pressure-containing internal members is at least about 10 percent ofthe total amount of upright surface area contacting the reaction mediumthat is attributable to the pressure-containing sidewalls of the vessel.More preferably, the total exposed upright surface area presented by theinternal members and contacting the reaction medium is in the range offrom about 15 to about 600 percent of the total exposed upright surfacearea presented by the pressure-containing sidewalls and contacting thereaction medium, still more preferably in the range of from about 25 toabout 400 percent, and most preferably in the range of from 35 to 200percent. The presence of added upright surface area in bubble columnoxidation can permit lower H:W ratios than would be possible inconventional bubble column reactors having little or no added uprightsurface area. Thus, it is preferred for the H:W ratio of a bubble columnreactor employing added upright surface area to be in the range of fromabout 3:1 to about 20:1, more preferably in the range of from about3.5:1 to about 15:1, and most preferably in the range of from 4:1 to12:1

Referring now to the embodiment illustrated in FIGS. 33 and 34,oxidation bubble column reactor 20 can include a single divider wall 33extending diametrically from one side of sidewall 46 to the oppositeside of sidewall 46. Divider wall 33 is disposed above sparger 34 sothat substantially all of the oxidant stream is introduced below thebottom of divider wall 33. Thus, approximately one-half of the gas phaseof reaction medium 36, which includes an undissolved portion of theoxidant stream, flows upwardly on each side of divider wall 33.Preferably, approximately one-half of the feed stream is introduced viafeed inlet 32 a on one side of divider wall 33, and the other half ofthe feed stream is introduced via feed inlet 32 b on the other side ofdivider wall 33. The height of divider wall 33 is substantially equal tothe height of cylindrical sidewall 46.

Divider wall 33 divides reaction zone 28 into approximately two halveswith reaction medium 36 being disposed on each side of divider wall 33.Substantially all of the upright surface area of reactor 20 thatcontacts reaction medium 36 is attributable to the inner exposedsurfaces of sidewall 46 and the outer exposed surfaces of divider wall33. If divider wall 33 were not present in reaction zone 28, thensubstantially all of the upright surface area in contact with reactionmedium 36 would be attributable to sidewalls 46 of thepressure-containing vessel shell 22. Divider wall 33 provides additionalsurface area that affects the fluid flow dynamics of reaction medium 36and permits oxidation bubble column reactor 20 to be scaled-up withoutsignificant negative effects on reactor performance.

Referring now to the embodiment illustrated in FIG. 35, oxidation bubblecolumn reactor 20 is illustrated as including a truncated divider wall35. Divider wall 35 of FIG. 35 is similar to divider wall 33 of FIGS. 33and 34; however, divider wall 35 of FIG. 35 has a height that issignificantly less than the overall height of reaction medium 36. In theconfiguration illustrated in FIG. 35, is preferred for substantially allof the feed stream and the oxidant stream to be introduced into reactionzone 28 below the bottom of divider wall 35. The top of divider wall 35is preferably spaced a substantial distance from the upper surface 44 ofreaction medium 36. In this configuration the two halves of reactionmedium 36 disposed on either side of divider wall 35 can mix with oneanother above and below divider wall 35.

Referring now to the embodiment illustrated in FIGS. 36 and 37,oxidation bubble column reactor 20 can include a non-planer divider wall37 which permits a substantial amount of upright surface area to beadded to reactor 20, without requiring a plurality of additionalinternal members in reaction zone 28. Like divider walls 33,35 of FIGS.33-35, divider wall 37 of FIGS. 36 and 37 is coupled to and extendsbetween diametrically opposed inner sidewall surfaces of sidewall 46.

In the embodiment illustrated in FIGS. 38 and 39, oxidation bubblecolumn reactor 20 includes an upright internal member 41 having agenerally X-shaped configuration. The outer vertical edges of internalmember 41 are spaced inwardly from the inner surfaces of sidewall 46 sothat reaction medium 36 can flow between the partial quadrants definedby X-shaped internal member 41. The various exposed outer uprightsurfaces of internal member 71 add a significant amount of surface areathat contacts reaction medium 36.

FIGS. 40-42 illustrate an embodiment where a portion of reaction zone 28is divided into 4 vertical quadrants via internal member 53, whileanother portion of reaction zone 28 is divided into 8 verticalwedge-shaped sections via internal member 55. As illustrated in FIG. 40,reaction zone 28 alternates vertically between division into 4 verticalquadrants with internal member 53 and 8 vertical wedge-shaped sectionsvia internal member 55.

Referring now to FIGS. 43 and 44, oxidation bubble column reactor 20 isillustrated as including a plurality of generally helicoid-shapedinternal members 61 a,b,c,d and a vertical X-shaped divider member 63disposed above helical members 61. Helicoid members 61 present slopedouter surfaces that induce a swirling flow pattern in the upwardlyflowing portion of reaction medium 36. It is preferred for the directionof slope of helicoid members 61 to be such that adjacent helical members61 cause swirling of reaction medium 36 in generally oppositedirections. Thus, if helicoid member 61 a causes reaction medium torotate clockwise as reaction medium 36 rises in reaction zone 28,helicoid member 61 b (disposed immediately above helical member 61 a)causes the upwardly moving reaction medium 36 to swirl in acounter-clockwise fashion. Vertical internal divider member 93 addsadditional upright surface area to oxidation bubble column reactor 20,and can also function to minimize swirling/turbulence of reaction medium36 as the gas-phase rises towards upper surface 44 of reaction medium36.

Regardless of which configuration illustrated in FIGS. 33-44 is employedin oxidation bubble column reactor 20, it is preferred for the oxidantstream and the feed stream to be introduced into reaction zone 28 in amanner such that a substantial portion of the molecular oxygen andoxidizable compound are introduced into reaction zone 28 below asignificant portion of the upright internal member or members.Preferably, at least about 50 weight percent of all the molecular oxygenand oxidizable compound introduced into reaction zone 28 enters reactionzone 28 below at least 50 percent of the upright exposed outer surfacearea of the internal member or members, more preferably at least about75 weight percent of the molecular oxygen and oxidizable compound enterreaction zone below at least 50 percent of the upright exposed outersurface area of the internal member(s), still more preferably at leastabout 90 weight percent of the molecular oxygen and oxidizable compoundenter reaction zone below at least 50 percent of the upright exposedouter surface area of the internal member(s), and most preferablysubstantially all of the molecular oxygen and oxidizable compound enterreaction zone below at least 50 percent of the upright exposed outersurface area of the internal member(s). In addition, it is preferred forthe oxidant stream and the feed stream to be introduced in a manner suchthat a substantial portion of the gas phase of reaction medium 36 flowsupwardly on all sides of the additional exposed outer surface areaprovided by the internal member or members. It is further preferred forthe oxidant and feed streams to be introduced into reaction zone 28 inaccordance with the radial and azimuthal distribution schemes describedabove.

Although certain prior art oxidation reactors may employ heat exchangesurfaces that contact the reaction medium in a manner similar to that ofthe internal member(s) described herein, it should be noted that it isundesirable for the internal member(s) of the present invention toprovide any significant degree of heating or cooling to the reactionmedium. Thus, it is preferred for the heat flux of the exposed (i.e.,contacting the reaction medium) upright surfaces of the internalmember(s) described herein to be less than about 30,000 watts per squaremeter.

FIGS. 45-53 illustrate embodiments of the present invention whereoxidation bubble column reactor 20 is equipped with one or more bafflesthat contact multi-phase reaction medium 36 in order to facilitateimproved oxidation with minimal impurities formation. Baffles areespecially useful in the scaled-up bubble column reactor designsdescribed above. In each of the baffled bubble column reactors 20illustrated in FIGS. 45-53, it is preferred for the baffle or baffles tohave an open area in the range of from about 10 to about 90 percent. Asused herein, percent open area of a baffle means the minimum percent ofthe horizontal cross-sectional area of a reaction zone that is open(i.e., not filled by the structure of the baffle) at the verticallocation of the baffle. More preferably, the open area of the baffle orbaffles illustrated in FIGS. 45-53 is in the range of from about 25 toabout 75 percent, most preferably in the range of from 35 to 65 percent.

A significant feature of the baffles illustrated in FIGS. 45-53 is theability of the baffles to resist fouling. As mentioned previously,oxidation bubble column reactor 20 is preferably employed inprecipitating oxidation service, where solids are formed in reactionmedium 36 during oxidation. Baffles having a significant amount ofnear-horizontal upwardly-facing planar surface area are prone to foulingin a reactor operating under precipitating conditions. When bafflesfoul, solids build up on the upwardly-facing surfaces of the baffles,and as the amount of solids deposited on the baffles increases, chunksof the precipitated solids may dislodge from the baffles and falltowards the bottom of the reactor. These chunks of dislodged solids canbuild up in the bottom of the reactor and can cause a number of problemsincluding, for example, inhibition of slurry discharge out of the bottomof the reactor.

In view of the foregoing, it is preferred in one embodiment for thebaffle or baffles employed in oxidation bubble column reactor 20 topresent no upwardly-facing planar outer surfaces (e.g., the baffle canbe constructed from piping materials having a circular cross section).Unless otherwise defined herein, an upwardly-facing surface is a surfacehaving a normal vector projecting above horizontal. In anotherembodiment, a small amount of substantially planar surfaces can beutilized so long as less than about 50 percent of the totalupwardly-facing exposed (i.e., contacting reaction medium 36) outersurface area of the baffle or baffles is attributable to substantiallyplanar surfaces inclined less than 30°, or 20°, or even 10° fromhorizontal. More preferably, less than about 35 percent of the totalupwardly-facing exposed outer surface area of the baffle or baffles isattributable to planar surfaces inclined less than 30°, or 20°, or even10° from horizontal. Most preferably, less than 25 percent of the totalupwardly-facing exposed outer surface area of the baffle or baffles isattributable to substantially planar surfaces inclined less than 30°, or20°, or even 10° from horizontal. It is further preferred for theupwardly-facing exposed outer surfaces of the baffle or baffles to havea substantially smooth finish so as to further resist fouling.Preferably, at least a substantial portion of the upwardly-facingexposed outer surfaces of the baffle or baffles have a surface roughnessless than about 125 micron RMS, more preferably less than about 64micron RMS, and most preferably less than 32 micron RMS.Electro-polished finishes and smooth “2 B” mill rolled finishes areparticularly useful.

In addition to the non-fouling design of the baffle or bafflesillustrated in FIGS. 45-53, is preferred for the baffle(s) to beproperly spaced from the upper and lower ends of reaction medium 36 soas to provide optimized effectiveness. In a preferred embodiment of thepresent invention the baffle or baffles are spaced at least 0.5 W and/or0.05 H from both the upper and lower ends of reaction medium 36, where“W” is the maximum width of reaction medium 36 and “H” is the maximumheight of reaction medium 36. More preferably, the baffle or baffles arespaced at least 1 W and/or 0.1 H from both the upper and lower ends ofreaction medium 36. Most preferably, the baffle or baffles are spaced atleast 1.5 W and/or 0.15 H from both the upper and lower ends of reactionmedium 36. The presence of a baffle or baffles in oxidation bubblecolumn reactor 20 can permit lower H:W ratios than would be possible insimilar non-baffled reactors. Thus, it is preferred for the H:W ratio ofbaffled bubble column reactors to be in the range of from about 3:1 toabout 20:1, more preferably in the range of from about 3.5:1 to about15:1, and most preferably in the range of from 4:1 to 12:1.

Referring now to FIGS. 45-47 in detail, oxidation bubble column reactor20 is illustrated as including a plurality of vertically-spaced baffles71 a,b,c,d. Preferably, oxidation bubble column reactor 20 includes inthe range of from 2 to 20 vertically-spaced baffles, most preferably 3to 8 vertically-spaced baffles. Preferably, each baffle 71 comprises aplurality of elongated individual baffle members 73. In this embodiment,each individual baffle member 73 presents a substantially cylindricalexposed outer surface that contacts reaction medium 36. In theembodiment illustrated in FIGS. 45-47, baffles 71 a,b,c,d are rotatedrelative to one another so that individual baffle members 73 of adjacentbaffles 71 extend substantially perpendicular to one another.

Referring now to FIGS. 48-50 in detail, alternative baffles 81 a,b,c,dare illustrated as comprising a plurality of elongated individual bafflemembers 83. In this embodiment, each baffle member 83 is formed of anL-section member and presents a generally inverted V-shapedupwardly-facing exposed outer surface. The configuration of the exposedouter surfaces of individual baffle members 83 helps prevent foulingwith the precipitating solids in reaction medium 36. The number,spacing, and orientation of angle iron baffle members 83 can besubstantially the same as described above for cylindrical baffle members73 of FIGS. 45-47.

Referring now to FIGS. 51-53 in detail, oxidation bubble column reactor20 is illustrated as including a single monolithic baffle 91 havinggenerally the shape of two oppositely-extending vertical cones 93 a,bjoined at their base. The slope of the exposed upwardly-facing outersurface of monolithic baffle 91 helps prevent fouling of baffle 91 withthe solids precipitating in reaction medium 36.

The various baffle configurations illustrated in FIGS. 45-53 are onlyexemplary, and many other baffle configurations may fall within theambit of the present invention. It should also be noted that the baffleconfigurations illustrated in FIGS. 45-53 can be used in combination.

Although certain prior art oxidation reactors may employ heat exchangetubes that contact the reaction medium in a manner similar to that ofthe baffle(s) described herein, it should be noted that it isundesirable for the baffles of the present invention to provide anysignificant degree of heating or cooling of the reaction medium. Thus,it is preferred for the heat flux of the exposed (i.e., contacting thereaction medium) outer surfaces of the baffles described herein to beless than about 30,000 watts per square meter.

Referring again to FIGS. 1-53, oxidation is preferably carried out inbubble column reactor 20 under conditions that are markedly different,according to preferred embodiments disclosed herein, than conventionaloxidation reactors. When bubble column reactor 20 is used to carry outthe liquid-phase partial oxidation of para-xylene to crude terephthalicacid (CTA) according to preferred embodiments disclosed herein, thespatial profiles of local reaction intensity, of local evaporationintensity, and of local temperature combined with the liquid flowpatterns within the reaction medium and the preferred, relatively lowoxidation temperatures contribute to the formation of CTA particleshaving unique and advantageous properties.

FIGS. 54A and 54B illustrate base CTA particles produced in accordancewith one embodiment of the present invention. FIG. 54A shows the baseCTA particles at 500 times magnification, while FIG. 54B zooms in on oneof the base CTA particles and shows that particle at 2,000 timesmagnification. As perhaps best illustrated in FIG. 54B, each base CTAparticle is typically formed of a large number of small, agglomeratedCTA subparticles, thereby giving the base CTA particle a relatively highsurface area, high porosity, low density, and good dissolvability. Thebase CTA particles typically have a mean particle size in the range offrom about 20 to about 150 microns, more preferably in the range of fromabout 30 to about 120 microns, and most preferably in the range of from40 to 90 microns. The CTA subparticles typically have a mean particlesize in the range of from about 0.5 to about 30 microns, more preferablyfrom about 1 to about 15 microns, and most preferably in the range offrom 2 to 5 microns. The relatively high surface area of the base CTAparticles illustrated in FIGS. 54A and 54B, can be quantified using aBraunauer-Emmett-Teller (BET) surface area measurement method.Preferably, the base CTA particles have an average BET surface of atleast about 0.6 meters squared per gram (m²/g). More preferably, thebase CTA particles have an average BET surface area in the range of fromabout 0.8 to about 4 m²/g. Most preferably, the base CTA particles havean average BET surface area in the range of from 0.9 to 2 m²/g. Thephysical properties (e.g., particle size, BET surface area, porosity,and dissolvability) of the base CTA particles formed by optimizedoxidation process of a preferred embodiment of the present inventionpermit purification of the CTA particles by more effective and/oreconomical methods, as described in further detail below with respect toFIG. 57.

The mean particle size values provided above were determined usingpolarized light microscopy and image analysis. The equipment employed inthe particle size analysis included a Nikon E800 optical microscope witha 4x Plan Flour N.A. 0.13 objective, a Spot RT™ digital camera, and apersonal computer running Image Pro Plus™ V4.5.0.19 image analysissoftware. The particle size analysis method included the following mainsteps: (1) dispersing the CTA powders in mineral oil; (2) preparing amicroscope slide/cover slip of the dispersion; (3) examining the slideusing polarized light microscopy (crossed polars condition—particlesappear as bright objects on black background); (4) capturing differentimages for each sample preparation (field size=3×2.25 mm; pixelsize=1.84 microns/pixel); (5) performing image analysis with Image ProPlus™ software; (6) exporting the particle measures to a spreadsheet;and (7) performing statistical characterization in the spreadsheet. Step(5) of “performing image analysis with Image Pro Plus™ software”included the substeps of: (a) setting the image threshold to detectwhite particles on dark background; (b) creating a binary image; (c)running a single-pass open filter to filter out pixel noise; (d)measuring all particles in the image; and (e) reporting the meandiameter measured for each particle. The Image Pro Plus™ softwaredefines mean diameter of individual particles as the number averagelength of diameters of a particle measured at 2 degree intervals andpassing through the particle's centroid. Step 7 of “performingstatistical characterization in the spreadsheet” comprises calculatingthe volume-weighted mean particle size as follows. The volume of each ofthe n particles in a sample is calculated as if it were spherical usingpi/6*d_(i)ˆ3; multiplying the volume of each particle times its diameterto find pi/6*d_(i)ˆ4; summing for all particles in the sample of thevalues of pi/6*d_(i)ˆ4; summing the volumes of all particles in thesample; and calculating the volume-weighted particle diameter as sum forall n particles in the sample of (pi/6*d_(i)ˆ4) divided by sum for all nparticles in the sample of (pi/6*d_(i)ˆ3). As used herein, “meanparticle size” refers to the volume-weighted mean particle sizedetermined according to the above-described test method; and it is alsoreferred to as D(4,3).${D\left( {4,3} \right)} = \frac{\sum\limits_{i = 1}^{n}{\frac{\pi}{6}d_{i}^{4}}}{\sum\limits_{i = 1}^{n}{\frac{\pi}{6}d_{i}^{3}}}$

In addition, step 7 comprises finding the particle sizes for whichvarious fractions of the total sample volume are smaller. For example,D(v,0.1) is the particle size for which 10 percent of the total samplevolume is smaller and 90 percent is larger; D(v,0.5) is the particlesize for which one-half of the sample volume is larger and one-half issmaller; D(v,0.9) is the particle size for which 90 percent of the totalsample volume is smaller; and so on. In addition, step 7 comprisescalculating the value of D(v,0.9) minus D(v,0.1), which is hereindefined as the “particle size spread”; and step 7 comprises calculatingthe value of the particle size spread divided by D(4,3), which is hereindefined as the “particle size relative spread.”

Furthermore, it is preferable that the D(v,0.1) of the CTA particles asmeasured above be in the range from about 5 to about 65 microns, morepreferably in the range from about 15 to about 55 microns and mostpreferably in the range from 25 to 45 microns. It is preferable that theD(v,0.5) of the CTA particles as measured above be in the range fromabout 10 to about 90 microns, more preferably in the range from about 20to about 80 microns, and most preferably in the range from 30 to 70microns. It is preferable that the D(v,0.9) of the CTA particles asmeasured above be in the range from about 30 to about 150 microns, morepreferably in the range from about 40 to about 130 microns, and mostpreferably in the range from 50 to 110 microns. It is preferable thatthe particle size relative spread be in the range from about 0.5 toabout 2.0, more preferably in the range from about 0.6 to about 1.5, andmost preferably in the range from 0.7 to 1.3.

The BET surface area values provided above were measured on aMicromeritics ASAP2000 (available from Micromeritics InstrumentCorporation of Norcross, Ga.). In the first step of the measurementprocess, a 2 to 4 gram of sample of the particles was weighed and driedunder vacuum at 50° C. The sample was then placed on the analysis gasmanifold and cooled to 77°K. A nitrogen adsorption isotherm was measuredat a minimum of 5 equilibrium pressures by exposing the sample to knownvolumes of nitrogen gas and measuring the pressure decline. Theequilibrium pressures were appropriately in the range of P/P₀=0.01-0.20,where P is equilibrium pressure and P₀ is vapor pressure of liquidnitrogen at 77° K. The resulting isotherm was then plotted according tothe following BET equation:$\frac{P}{V_{a}\left( {P_{o} - P} \right)} = {\frac{1}{V_{m}C} + {\frac{C - 1}{V_{m}C}\left( \frac{P}{P_{o}} \right)}}$where V_(a) is volume of gas adsorbed by sample at P, V_(m) is volume ofgas required to cover the entire surface of the sample with a monolayerof gas, and C is a constant. From this plot, V_(m) and C weredetermined. V_(m) was then converted to a surface area using the crosssectional area of nitrogen at 77° K by:$A = {\sigma\quad\frac{V_{m}}{RT}}$where σ is cross sectional area of nitrogen at 77° K, T is 77° K, and Ris the gas constant.

As alluded to above, CTA formed in accordance with one embodiment of thepresent invention exhibits superior dissolution properties versesconventional CTA made by other processes. This enhanced dissolution rateallows the inventive CTA to be purified by more efficient and/or moreeffective purification processes. The following description addressesthe manner in which the rate of dissolution of CTA can quantified.

The rate of dissolution of a known amount of solids into a known amountof solvent in an agitated mixture can be measured by various protocols.As used herein, a measurement method called the “timed dissolution test”is defined as follows. An ambient pressure of about 0.1 megapascal isused throughout the timed dissolution test. The ambient temperature usedthroughout the timed dissolution test is about 22° C. Furthermore, thesolids, solvent and all dissolution apparatus are fully equilibratedthermally at this temperature before beginning testing, and there is noappreciable heating or cooling of the beaker or its contents during thedissolution time period. A solvent portion of fresh, HPLC analyticalgrade of tetrahydrofuran (>99.9 percent purity), hereafter THF,measuring 250 grams is placed into a cleaned KIMAX tall form 400milliliter glass beaker (Kimble® part number 14020, Kimble/Kontes,Vineland, N.J.), which is uninsulated, smooth-sided, and generallycylindrical in form. A Teflon-coated magnetic stirring bar (VWR partnumber 58948-230, about 1-inch long with ⅜-inch diameter, octagonalcross section, VWR International, West Chester, Pa. 19380) is placed inthe beaker, where it naturally settles to the bottom. The sample isstirred using a Variomag® multipoint 15 magnetic stirrer (H&PLabortechnik AG, Oberschleissheim, Germany) magnetic stirrer at asetting of 800 revolutions per minute. This stirring begins no more than5 minutes before the addition of solids and continues steadily for atleast 30 minutes after adding the solids. A solid sample of crude orpurified TPA particulates amounting to 250 milligrams is weighed into anon-sticking sample weighing pan. At a starting time designated as t=0,the weighed solids are poured all at once into the stirred THF, and atimer is started simultaneously. Properly done, the THF very rapidlywets the solids and forms a dilute, well-agitated slurry within 5seconds. Subsequently, samples of this mixture are obtained at thefollowing times, measured in minutes from t=0: 0.08, 0.25, 0.50, 0.75,1.00, 1.50, 2.00, 2.50, 3.00, 4.00, 5.00, 6.00, 8.00, 10.00, 15.00, and30.00. Each small sample is withdrawn from the dilute, well-agitatedmixture using a new, disposable syringe (Becton, Dickinson and Co, 5milliliter, REF 30163, Franklin Lakes, N.J. 07417). Immediately uponwithdrawal from the beaker, approximately 2 milliliters of clear liquidsample is rapidly discharged through a new, unused syringe filter (25 mmdiameter, 0.45 micron, Gelman GHP Acrodisc GF®, Pall Corporation, EastHills, N.Y. 11548) into a new, labeled glass sample vial. The durationof each syringe filling, filter placement, and discharging into a samplevial is correctly less than about 5 seconds, and this interval isappropriately started and ended within about 3 seconds either side ofeach target sampling time. Within about five minutes of each filling,the sample vials are capped shut and maintained at approximatelyconstant temperature until performing the following chemical analysis.After the final sample is taken at a time of 30 minutes past t=0, allsixteen samples are analyzed for the amount of dissolved TPA using aHPLC-DAD method generally as described elsewhere within this disclosure.However, in the present test, the calibration standards and the resultsreported are both based upon milligrams of dissolved TPA per gram of THFsolvent (hereafter “ppm in THF”). For example, if all of the 250milligrams of solids were very pure TPA and if this entire amount fullydissolved in the 250 grams of THF solvent before a particular samplewere taken, the correctly measured concentration would be about 1,000ppm in THF.

When CTA according to the present invention is subjected to the timeddissolution test described above, it is preferred that a sample taken atone minute past t=0 dissolves to a concentration of at least about 500ppm in THF, more preferably to at least 600 ppm in THF. For a sampletaken at two minutes past t=0, it is preferred that CTA according to thecurrent invention will dissolve to a concentration of at least about 700ppm in THF, more preferably to at least 750 ppm in THF. For a sampletaken at four minutes past t=0, it is preferred that CTA according tothe current invention will dissolve to a concentration of at least about840 ppm in THF, more preferably to at least 880 ppm in THF.

The inventors have found that a relatively simple negative exponentialgrowth model is useful to describe the time dependence of the entiredata set from a complete timed dissolution test, notwithstanding thecomplexity of the particulate samples and of the dissolution process.The form of the equation, hereinafter the “timed dissolution model”, isas follows:S=A+B*(1−exp(−C*t)), where

-   -   t=time in units of minutes;    -   S=solubility, in units of ppm in THF, at time t;    -   exp=exponential function in the base of the natural logarithm of        2;    -   A, B=regressed constants in units of ppm in THF, where A relates        mostly to the rapid dissolution of the smaller particles at very        short times, and where the sum of A+B relates mostly to the        total amount of dissolution near the end of the specified        testing period; and    -   C=a regressed time constant in units of reciprocal minutes.

The regressed constants are adjusted to minimize the sum of the squaresof the errors between the actual data points and the corresponding modelvalues, which method is commonly called a “least squares” fit. Apreferred software package for executing this data regression is JMPRelease 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary,N.C. 27513).

When CTA according to the present invention is tested with the timeddissolution test and fitted to the timed dissolution model describedabove, it is preferred for the CTA to have a time constant “C” greaterthan about 0.5 reciprocal minutes, more preferably greater than about0.6 reciprocal minutes, and most preferably greater than 0.7 reciprocalminutes.

FIGS. 55A and 55B illustrate a conventional CTA particle made by aconventional high-temperature oxidation process in a continuous stirredtank reactor (CSTR). FIG. 55A shows the conventional CTA particle at 500times magnification, while FIG. 55B zooms in and shows the CTA particleat 2,000 times magnification. A visual comparison of the inventive CTAparticles illustrated in FIGS. 54A and 54B and the conventional CTAparticle illustrated in FIGS. 55A and 55B shows that the conventionalCTA particle has a higher density, lower surface area, lower porosity,and larger particle size than the inventive CTA particles. In fact, theconventional CTA represented in FIGS. 55A and 55B has a mean particlesize of about 205 microns and a BET surface area of about 0.57 m²/g.

FIG. 56 illustrates a conventional process for making purifiedterephthalic acid (PTA). In the conventional PTA process, para-xylene ispartially oxidized in a mechanically agitated high temperature oxidationreactor 700. A slurry comprising CTA is withdrawn from reactor 700 andthen purified in a purification system 702. The PTA product ofpurification system 702 is introduced into a separation system 706 forseparation and drying of the PTA particles. Purification system 702represents a large portion of the costs associated with producing PTAparticles by conventional methods. Purification system 702 generallyincludes a water addition/exchange system 708, a dissolution system 710,a hydrogenation system 712, and three separate crystallization vessels704 a,b,c. In water addition/exchange system 708, a substantial portionof the mother liquor is displaced with water. After water addition, thewater/CTA slurry is introduced into the dissolution system 710 where thewater/CTA mixture is heated until the CTA particles fully dissolve inthe water. After CTA dissolution, the CTA-in-water solution is subjectedto hydrogenation in hydrogenation system 712. The hydrogenated effluentfrom hydrogenation system 712 is then subjected to three crystallizationsteps in crystallization vessels 704 a,b,c, followed by PTA separationin separation system 706.

FIG. 57 illustrates an improved process for producing PTA employing abubble column oxidation reactor 800 configured in accordance with anembodiment of the present invention. An initial slurry comprising solidCTA particles and a liquid mother liquor is withdrawn from reactor 800.Typically, the initial slurry may contain in the range of from about 10to about 50 weight percent solid CTA particles, with the balance beingliquid mother liquor. The solid CTA particles present in the initialslurry typically contain at least about 400 ppmw of4-carboxybenzaldehyde (4-CBA), more typically at least about 800 ppmw of4-CBA, and most typically in the range of from 1,000 to 15,000 ppmw of4-CBA. The initial slurry withdrawn from reactor 800 is introduced intoa purification system 802 to reduce the concentration of 4-CBA and otherimpurities present in the CTA. A purer/purified slurry is produced frompurification system 802 and is subjected to separation and drying in aseparation system 804 to thereby produce purer solid terephthalic acidparticles comprising less than about 400 ppmw of 4-CBA, more preferablyless than about 250 ppmw of 4-CBA, and most preferably in the range offrom 10 to 200 ppmw of 4-CBA.

Purification system 802 of the PTA production system illustrated in FIG.57 provides a number of advantages over purification system 802 of theprior art system illustrated in FIG. 56. Preferably, purification system802 generally includes a liquor exchange system 806, a digester 808, anda single crystallizer 810. In liquor exchange system 806, at least about50 weight percent of the mother liquor present in the initial slurry isreplaced with a fresh replacement solvent to thereby provide asolvent-exchanged slurry comprising CTA particles and the replacementsolvent. The solvent-exchanged slurry exiting liquor exchange system 806is introduced into digester (or secondary oxidation reactor) 808. Indigester 808, a secondary oxidation reaction is preformed at slightlyhigher temperatures than were used in the initial/primary oxidationreaction carried out in bubble column reactor 800. As discussed above,the high surface area, small particle size, and low density of the CTAparticles produced in reactor 800 cause certain impurities trapped inthe CTA particles to become available for oxidation in digester 808without requiring complete dissolution of the CTA particles in digester808. Thus, the temperature in digester 808 can be lower than manysimilar prior art processes. The secondary oxidation carried out indigester 808 preferably reduces the concentration of 4-CBA in the CTA byat least 200 ppmw, more preferably at least about 400 ppmw, and mostpreferably in the range of from 600 to 6,000 ppmw. Preferably, thesecondary oxidation temperature in digester 808 is at least about 10° C.higher than the primary oxidation temperature in bubble column reactor800, more preferably about 20 to about 80° C. higher than the primaryoxidation temperature in reactor 800, and most preferably 30 to 50° C.higher than the primary oxidation temperature in reactor 800. Thesecondary oxidation temperature is preferably in the range of from about160 to about 240° C., more preferably in the range of from about 180 toabout 220° C. and most preferably in the range of from 190 to 210° C.The purified product from digester 808 requires only a singlecrystallization step in crystallizer 810 prior to separation inseparation system 804. Suitable secondary oxidation/digestion techniquesare discussed in further detail in U.S. patent application Pub. No.2005/0065373, the entire disclosure of which is expressly incorporatedherein by reference.

Terephthalic acid (e.g., PTA) produced by the system illustrated in FIG.57 is preferably formed of PTA particles having a mean particle size ofat least about 40 microns, more preferably in the range of from about 50to about 2,000 microns, and most preferably in the range of from 60 to200 microns. The PTA particles preferably have an average BET surfacearea less than about 0.25 m²/g, more preferably in the range of fromabout 0.005 to about 0.2 m²/g, and most preferably in the range of from0.01 to 0.18 m²/g. PTA produced by the system illustrated in FIG. 57 issuitable for use as a feedstock in the making of PET. Typically, PET ismade via esterification of terephthalic with ethylene glycol, followedby polycondensation. Preferably, terephthalic acid produced by anembodiment of the present invention is employed as a feed to the pipereactor PET process described in U.S. patent application Ser. No.10/013,318, filed Dec. 7, 2001, the entire disclosure of which isincorporated herein by reference.

CTA particles with the preferred morphology disclosed herein areparticularly useful in the above-described oxidative digestion processfor reduction of 4-CBA content. In addition, these preferred CTAparticles provide advantages in a wide range of other post-processesinvolving dissolution and/or chemical reaction of the particles. Theseadditional post-processes include, but are not limited too, reactionwith at least one hydroxyl-containing compound to form ester compounds,especially the reaction of CTA with methanol to form dimethylterephthalate and impurity esters; reaction with at least one diol toform ester monomer and/or polymer compounds, especially the reaction ofCTA with ethylene glycol to form polyethylene terephthalate (PET); andfull or partial dissolution in solvents, including, but not limited too,water, acetic acid, and N-methyl-2-pyrrolidone, which may includefurther processing, including, but not limited too, reprecipitation of amore pure terephthalic acid and/or selective chemical reduction ofcarbonyl groups other than carboxylic acid groups. Notably included isthe substantial dissolution of CTA in a solvent comprising water coupledwith partial hydrogenation that reduces the amount of aldehydes,especially 4-CBA, fluorenones, phenones, and/or anthraquinones.

The inventors also contemplate that CTA particles having the preferredproperties disclosed herein can be produced from CTA particles notconforming to the preferred properties disclosed herein (non-conformingCTA particles) by means including, but not limited too, mechanicalcomminution of non-conforming CTA particles and full or partialdissolution of non-conforming CTA particles followed by full or partialre-precipitation.

In accordance with one embodiment of the present invention, there isprovided a process for partially oxidizing an oxidizable aromaticcompound to one or more types of aromatic carboxylic acid wherein thepurity of the solvent portion of the feed (i.e., the “solvent feed”) andthe purity of the oxidizable compound portion of the feed (i.e., the“oxidizable compound feed”) are controlled within certain rangesspecified below. Along with other embodiments of the present invention,this enables the purity of the liquid phase and, if present, the solidphase and the combined slurry (i.e., solid plus liquid) phase of thereaction medium to be controlled in certain preferred ranges, outlinedbelow.

With respect to the solvent feed, it is known to oxidize an oxidizablearomatic compound(s) to produce an aromatic carboxylic acid wherein thesolvent feed introduced into the reaction medium is a mixture ofanalytical-purity acetic acid and water, as is often employed atlaboratory scale and pilot scale. Likewise, it is known to conduct theoxidation of oxidizable aromatic compound to aromatic carboxylic acidwherein the solvent leaving the reaction medium is separated from theproduced aromatic carboxylic acid and then recycled back to the reactionmedium as feed solvent, primarily for reasons of manufacturing cost.This solvent recycling causes certain feed impurities and processby-products to accumulate over time in the recycled solvent. Variousmeans are known in the art to help purify recycled solvent beforere-introduction into the reaction medium. Generally, a higher degree ofpurification of the recycled solvent leads to significantly highermanufacturing cost than does a lower degree of purification by similarmeans. One embodiment of the present invention relates to understandingand defining the preferred ranges of a large number of impurities withinthe solvent feed, many of which were heretofore thought largely benign,in order to find an optimal balance between overall manufacturing costand overall product purity.

“Recycled solvent feed” is defined herein as solvent feed comprising atleast about 5 weight percent mass that has previously passed through areaction medium containing one or more oxidizable aromatic compoundsundergoing partial oxidation. For reasons of solvent inventory and ofon-stream time in a manufacturing unit, it is preferable that portionsof recycled solvent pass through reaction medium at least once per dayof operation, more preferably at least once per day for at least sevenconsecutive days of operation, and most preferably at least once per dayfor at least 30 consecutive days of operation. For economic reasons, itis preferable that at least about 20 weight percent of the solvent feedto the reaction medium of the present invention is recycled solvent,more preferably at least about 40 weight percent, still more preferablyat least about 80 weight percent, and most preferably at least 90 weightpercent.

The inventors have discovered that, for reasons of reaction activity andfor consideration of metallic impurities left in the oxidation product,the concentrations of selected multivalent metals within the recycledsolvent feed are preferably in ranges specified immediately below. Theconcentration of iron in recycled solvent is preferably below about 150ppmw, more preferably below about 40 ppmw, and most preferably between 0and 8 ppmw. The concentration of nickel in recycled solvent ispreferably below about 150 ppmw, more preferably below about 40 ppmw,and most preferably between 0 and 8 ppmw. The concentration of chromiumin recycled solvent is preferably below about 150 ppmw, more preferablybelow about 40 ppmw, and most preferably between 0 and 8 ppmw. Theconcentration of molybdenum in recycled solvent is preferably belowabout 75 ppmw, more preferably below about 20 ppmw, and most preferablybetween 0 and 4 ppmw. The concentration of titanium in recycled solventis preferably below about 75 ppmw, more preferably below about 20 ppmw,and most preferably between 0 and 4 ppmw. The concentration of copper inrecycled solvent is preferably below about 20 ppmw, more preferablybelow about 4 ppmw, and most preferably between 0 and 1 ppmw. Othermetallic impurities are also typically present in recycled solvent,generally varying at lower levels in proportion to one or more of theabove listed metals. Controlling the above listed metals in thepreferred ranges will keep other metallic impurities at suitable levels.

These metals can arise as impurities in any of the incoming processfeeds (e.g., in incoming oxidizable compound, solvent, oxidant, andcatalyst compounds). Alternatively, the metals can arise as corrosionproducts from any of the process units contacting reaction medium and/orcontacting recycled solvent. The means for controlling the metals in thedisclosed concentration ranges include the appropriate specification andmonitoring of the purity of various feeds and the appropriate usage ofmaterials of construction, including, but not limited to, manycommercial grades of titanium and of stainless steels including thosegrades known as duplex stainless steels and high molybdenum stainlesssteels.

The inventors have also discovered preferred ranges for selectedaromatic compounds in the recycled solvent. These include bothprecipitated and dissolved aromatic compounds within the recycledsolvent.

Surprisingly, even precipitated product (e.g., TPA) from a partialoxidation of para-xylene, is a contaminant to be managed in recycledsolvent. Because there are surprisingly preferred ranges for the levelsof solids within the reaction medium, any precipitated product in thesolvent feed directly subtracts from the amount of oxidizable compoundthat can be fed in concert. Furthermore, feeding precipitated TPA solidsin the recycled solvent at elevated levels has been discovered to affectadversely the character of the particles formed within a precipitatingoxidation medium, leading to undesirable character in downstreamoperations (e.g., product filtration, solvent washing, oxidativedigestion of crude product, dissolution of crude product for furtherprocessing, and so on). Another undesirable characteristic ofprecipitated solids in the recycle solvent feed is that these oftencontain very high levels of precipitated impurities, as compared toimpurity concentrations in the bulk of the solids within the TPAslurries from which much of the recycled solvent is obtained. Possibly,the elevated levels of impurities observed in solids suspended inrecycled filtrate may relate to nucleation times for precipitation ofcertain impurities from the recycled solvent and/or to cooling of therecycled solvent, whether intentional or due to ambient losses. Forexample, concentrations of highly-colored and undesirable2,6-dicarboxyfluorenone have been observed at far higher levels insolids present in recycled solvent at 80° C. than are observed in TPAsolids separated from recycled solvent at 160° C. Similarly,concentrations of isophthalic acid have been observed at much higherlevels in solids present in recycled solvent compared to levels observedin TPA solids from the reaction medium. Exactly how specificprecipitated impurities entrained within recycled solvent behave whenre-introduced to the reaction medium appears to vary. This dependsperhaps upon the relative solubility of the impurity within the liquidphase of the reaction medium, perhaps upon how the precipitated impurityis layered within the precipitated solids, and perhaps upon the localrate of TPA precipitation where the solid first re-enters the reactionmedium. Thus, the inventors have found it useful to control the level ofcertain impurities in the recycled solvent, as disclosed below, withoutrespect to whether these impurities are present in the recycled solventin dissolved form or are entrained particulates therein.

The amount of precipitated solids present in recycled filtrate isdetermined by a gravimetric method as follows. A representative sampleis withdrawn from the solvent supply to the reaction medium while thesolvent is flowing in a conduit toward the reaction medium. A usefulsample size is about 100 grams captured in a glass container havingabout 250 milliliters of internal volume. Before being released toatmospheric pressure, but while continuously flowing toward the samplecontainer, the recycled filtrate is cooled to less than 100° C.; thiscooling is in order to limit solvent evaporation during the shortinterval before being sealed closed in the glass container. After thesample is captured at atmospheric pressure, the glass container issealed closed immediately. Then the sample is allowed to cool to about20° C. while surrounded by air at about 20° C. and without forcedconvection. After reaching about 20° C., the sample is held at thiscondition for at least about 2 hours. Then, the sealed container isshaken vigorously until a visibly uniform distribution of solids isobtained. Immediately thereafter, a magnetic stirrer bar is added to thesample container and rotated at sufficient speed to maintain effectivelyuniform distribution of solids. A 10 milliliter aliquot of the mixedliquid with suspended solids is withdrawn by pipette and weighed. Thenthe bulk of the liquid phase from this aliquot is separated by vacuumfiltration, still at about 20° C. and effectively without loss ofsolids. The moist solids filtered from this aliquot are then dried,effectively without sublimation of solids, and these dried solids areweighed. The ratio of the weight of the dried solids to the weight ofthe original aliquot of slurry is the fraction of solids, typicallyexpressed as a percentage and referred to herein as the recycledfiltrate content of precipitated solids at 20° C.

The inventors have discovered that aromatic compounds dissolved in theliquid phase of the reaction medium and comprising aromatic carboxylicacids lacking non-aromatic hydrocarbyl groups (e.g., isophthalic acid,benzoic acid, phthalic acid, 2,5,4′-tricarboxybiphenyl) are surprisinglypernicious components. Although these compounds are much reduced inchemical activity in the subject reaction medium compared to oxidizablecompounds having non-aromatic hydrocarbyl groups, the inventors havediscovered that these compounds nonetheless undergo numerous detrimentalreactions. Thus, it is advantageous to control the content of thesecompounds in preferred ranges in the liquid phase of the reactionmedium. This leads to preferred ranges of select compounds in recycledsolvent feed and also to preferred ranges of select precursors in theoxidizable aromatic compound feed.

For example, in the liquid-phase partial oxidation of para-xylene toterephthalic acid (TPA), the inventors have discovered that thehighly-colored and undesirable impurity 2,7-dicarboxyfluorenone(2,7-DCF) is virtually undetectable in the reaction medium and productoff-take when meta-substituted aromatic compounds are at very low levelsin the reaction medium. The inventors have discovered that whenisophthalic acid impurity is present at increasing levels in the solventfeed, the formation of 2,7-DCF rises in almost direct proportion. Theinventors have also discovered that when meta-xylene impurity is presentin the feed of para-xylene, the formation of 2,7-DCF again rises almostin direct proportion. Furthermore, even if the solvent feed andoxidizable compound feed are devoid of meta-substituted aromaticcompounds, the inventors have discovered that some isophthalic acid isformed during a typical partial oxidation of very pure para-xylene,particularly when benzoic acid is present in the liquid phase of thereaction medium. This self-generated isophthalic acid may, owing to itsgreater solubility than TPA in solvent comprising acetic acid and water,build up over time in commercial units employing recycled solvent. Thus,the amount of isophthalic acid within solvent feed, the amount ofmeta-xylene within oxidizable aromatic compound feed, and the rate ofself-creation of isophthalic acid within the reaction medium are allappropriately considered in balance with each other and in balance withany reactions that consume isophthalic acid. Isophthalic acid has beendiscovered to undergo additional consumptive reactions besides theformation of 2,7-DCF, as are disclosed below. In addition, the inventorshave discovered that there are other issues to consider when settingappropriate ranges for the meta-substituted aromatic species in thepartial oxidation of para-xylene to TPA. Other highly-colored andundesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF),appear to relate greatly to dissolved, para-substituted aromaticspecies, which are always present with para-xylene feed to aliquid-phase oxidation. Thus, the suppression of 2,7-DCF is bestconsidered in perspective with the level of other colored impuritiesbeing produced.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that the formation of trimelliticacid rises as the levels isophthalic acid and phthalic acid rise withinthe reaction medium. Trimellitic acid is a tri-functional carboxylicacid leading to branching of polymer chains during production of PETfrom TPA. In many PET applications, branching levels must be controlledto low levels and hence trimellitic acid must be controlled to lowlevels in purified TPA. Besides leading to trimellitic acid, thepresence of meta-substituted and ortho-substituted species in thereaction medium also give rise to other tricarboxylic acids (e.g.,1,3,5-tricarboxybenzene). Furthermore, the increased presence oftricarboxylic acids in the reaction medium increases the amount oftetracarboxylic acid formation (e.g., 1,2,4,5-tetracarboxybenzene).Controlling the summed production of all aromatic carboxylic acidshaving more than two carboxylic acid groups is one factor in setting thepreferred levels of meta-substituted and ortho-substituted species inthe recycled solvent feed, in the oxidizable compound feed, and in thereaction medium according to the present invention.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that increased levels in the liquidphase of the reaction medium of several dissolved aromatic carboxylicacids lacking non-aromatic hydrocarbyl groups leads directly to theincreased production of carbon monoxide and carbon dioxide. Thisincreased production of carbon oxides represents a yield loss on bothoxidant and on oxidizable compound, the later since many of theco-produced aromatic carboxylic acids, which on the one hand may beviewed as impurities, on the other hand also have commercial value.Thus, appropriate removal of relatively soluble carboxylic acids lackingnon-aromatic hydrocarbyl groups from recycle solvent has an economicvalue in preventing yield loss of oxidizable aromatic compound and ofoxidant, in addition to suppressing the generation of highly undesirableimpurities such as various fluorenones and trimellitic acid.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that formation of2,5,4′-tricarboxybiphenyl is seemingly unavoidable. The2,5,4′-tricarboxybiphenyl is an aromatic tricarboxylic acid formed bythe coupling of two aromatic rings, perhaps by the coupling of adissolved para-substituted aromatic species with an aryl radical,perhaps an aryl radical formed by decarboxylation or decarbonylation ofa para-substituted aromatic species. Fortunately, the2,5,4′-tricarboxybiphenyl is typically produced at lower levels thantrimellitic acid and does not usually lead to significantly increaseddifficulties with branching of polymer molecules during production ofPET. However, the inventors have discovered that elevated levels of2,5,4′-tricarboxybiphenyl in a reaction medium comprising oxidation ofalkyl aromatics according to preferred embodiments of the presentinvention lead to increased levels of highly-colored and undesirable2,6-DCF. The increased 2,6-DCF is possibly created from the2,5,4′-tricarboxybiphenyl by ring closure with loss of a water molecule,though the exact reaction mechanism is not known with certainty. If2,5,4′-tricarboxybiphenyl, which is more soluble in solvent comprisingacetic acid and water than is TPA, is allowed to build up too highwithin recycled solvent, conversion rates to 2,6-DCF can becomeunacceptably large.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that aromatic carboxylic acidslacking non-aromatic hydrocarbyl groups (e.g., isophthalic acid)generally lead to mild suppression of the chemical activity of thereaction medium when present in the liquid phase at sufficientconcentration.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that precipitation is very oftennon-ideal (i.e. non-equilibrium) with respect to the relativeconcentrations of different chemical species in the solid phase and inthe liquid phase. Perhaps, this is because the precipitation rate isvery fast at the space-time reaction rates preferred herein, leading tonon-ideal co-precipitation of impurities, or even occlusion. Thus, whenit is desired to limit the concentration of certain impurities (e.g.,trimellitic acid and 2,6-DCF) within crude TPA, owing to theconfiguration of downstream unit operations, it is preferable to controltheir concentration in solvent feed as well as their generation ratewithin the reaction medium.

For example, the inventors have discovered that benzophenone compounds(e.g., 4,4′-dicarboxybenzophenone and 2,5,4′-tricarboxybenzophenone)made during partial oxidation of para-xylene, have undesirable effectsin a PET reaction medium even though benzophenone compounds are not ashighly colored in TPA per se as are fluorenones and anthraquinones.Accordingly, it is desirable to limit the presence of benzophenones andselect precursors in recycled solvent and in oxidizable compound feed.Furthermore, the inventors have discovered that the presence of elevatedlevels of benzoic acid, whether admitted in recycled solvent or formedwithin the reaction medium, leads to elevated rates of production of4,4′-dicarboxybenzophenone.

In review, the inventors have discovered and sufficiently quantified asurprising array of reactions for aromatic compounds lackingnon-aromatic hydrocarbyl groups that are present in the liquid-phasepartial oxidation of para-xylene to TPA. Recapping just the single caseof benzoic acid, the inventors have discovered that increased levels ofbenzoic acid in the reaction medium of certain embodiments of thepresent invention lead to greatly increased production of the highlycolored and undesirable 9-fluorenone-2-carboxylic acid, to greatlyincreased levels of 4,4′-dicarboxybiphenyl, to increased levels of4,4′-dicarboxybenzophenone, to a mild suppression of chemical activityof the intended oxidation of para-xylene, and to increased levels ofcarbon oxides and attendant yield losses. The inventors have discoveredthat increased levels of benzoic acid in the reaction medium also leadto increased production of isophthalic acid and phthalic acid, thelevels of which are desirably controlled in low ranges according tosimilar aspects of the current invention. The number and importance ofreactions involving benzoic acid are perhaps even more surprising sincesome recent inventors contemplate using benzoic acid in place of aceticacid as a primary component of solvent (See, e.g., U.S. Pat. No.6,562,997). Additionally, the present inventors have observed thatbenzoic acid is self-generated during oxidation of para-xylene at ratesthat are quite important relative to its formation from impurities, suchas toluene and ethylbenzene, commonly found in oxidizable compound feedcomprising commercial-purity para-xylene.

On the other hand, the inventors have discovered little value fromadditional regulation of recycled solvent composition in regard to thepresence of oxidizable aromatic compound and in regard to aromaticreaction intermediates that both retain non-aromatic hydrocarbyl groupsand are also relatively soluble in the recycled solvent. In general,these compounds are either fed to or created within the reaction mediumat rates substantially greater than their presence in recycled solvent;and the consumption rate of these compounds within the reaction mediumis great enough, retaining one or more non-aromatic hydrocarbyl groups,to limit appropriately their build-up within recycled solvent. Forexample, during partial oxidation of para-xylene in a multi-phasereaction medium, para-xylene evaporates to a limited extent along withlarge quantities of solvent. When this evaporated solvent exits thereactor as part of the off-gas and is condensed for recovery as recycledsolvent, a substantial portion of the evaporated para-xylene condensestherein as well. It is not necessary to limit the concentration of thispara-xylene in recycled solvent. For example, if solvent is separatedfrom solids upon slurry exiting a para-xylene oxidation reaction medium,this recovered solvent will contain a similar concentration of dissolvedpara-toluic acid to that present at the point of removal from thereaction medium. Although it may be important to limit the standingconcentration of para-toluic acid within the liquid phase of thereaction medium, see below, it is not necessary to regulate separatelythe para-toluic acid in this portion of recycled solvent owing to itsrelatively good solubility and to its low mass flow rate relative to thecreation of para-toluic acid within the reaction medium. Similarly, theinventors have discovered little reason to limit the concentrations inrecycled solvent of aromatic compounds with methyl substituents (e.g.toluic acids), aromatic aldehydes (e.g., terephthaldehyde), of aromaticcompounds with hydroxy-methyl substituents (e.g., 4-hydroxymethylbenzoicacid), and of brominated aromatic compounds retaining at least onenon-aromatic hydrocarbyl group (e.g., alpha-bromo-para-toluic acid)below those inherently found in the liquid phase exiting from thereaction medium occurring in the partial oxidation of xylene accordingto preferred embodiments of the present invention. Surprisingly, theinventors have also discovered that it is also not necessary to regulatein recycled solvent the concentration of selected phenols intrinsicallyproduced during partial oxidation of xylene, for these compounds arecreated and destroyed within the reaction medium at rates much greaterthan their presence in recycled solvent. For example, the inventors havediscovered that 4-hydroxybenzoic acid has relatively small effects onchemical activity in the preferred embodiments of the present inventionwhen co-fed at rates of over 2 grams of 4-hydroxybenzoic acid per 1kilogram of para-xylene, far higher than the natural presence inrecycled solvent, despite being reported by others as a significantpoison in similar reaction medium (See, e.g., W. Partenheimer, CatalysisToday 23 (1995) p. 81).

Thus, there are numerous reactions and numerous considerations insetting the preferred ranges of various aromatic impurities in thesolvent feed as now disclosed. These discoveries are stated in terms ofthe aggregated weight average composition of all solvent streams beingfed to the reaction medium during the course of a set time period,preferably one day, more preferably one hour, and most preferably oneminute. For example, if one solvent feed flows substantiallycontinuously with a composition of 40 ppmw of isophthalic acid at a flowrate of 7 kilograms per minute, a second solvent feed flowssubstantially continuously with a composition of 2,000 ppmw ofisophthalic acid at a flow rate of 10 kilograms per minute, and thereare no other solvent feed streams entering the reaction medium, then theaggregated weight average composition of the solvent feed is calculatedas (40*7+2,000*10)/(7+10)=1,193 ppmw of isophthalic acid. It is notablethat the weight of any oxidizable compound feed or of any oxidant feedthat are perhaps commingled with the solvent feed before entering thereaction medium are not considered in calculating the aggregated weightaverage composition of the solvent feed.

Table 1, below, lists preferred values for certain components in thesolvent feed introduced into the reaction medium. The solvent feedcomponents listed in Table 1 are as follows: 4-carboxybenzaldehyde(4-CBA), 4,4′-dicarboxystilbene (4,4′-DCS), 2,6-dicarboxyanthraquinone(2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone(2,7-DCF), 3,5-dicarboxyfluorenone (3,5DCF), 9-fluorenone-2-carboxylicacid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), totalfluorenones including other fluorenones not individually listed (totalfluorenones), 4,4′-dicarboxybiphenyl (4,4′-DCB),2,5,4′-tricarboxybiphenyl (2,5,4′-TCB), phthalic acid (PA), isophthalicacid (IPA), benzoic acid (BA), trimellitic acid (TMA),2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4′-dicarboxybenzil (4,4′-DCBZ),4,4′-dicarboxybenzophenone (4,4′-DCBP), 2,5,4′-tricarboxybenzophenone(2,5,4′-TCBP), terephthalic acid (TPA), precipitated solids at 20° C.,and total aromatic carboxylic acids lacking non-aromatic hydrocarbylgroups. Table 1, below provides the preferred amounts of theseimpurities in CTA produced according to an embodiment of the presentinvention. TABLE 1 Components of Solvent Feed Introduced into ReactionMedium Component Preferred More Preferred Most Preferred IdentificationAmt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA <1,200  30-600  60-3004,4′-DCS <3 <2 <1 2,6-DCA <6 0.1-3   0.2-1   2,6-DCF <20 0.1-10  0.5-5  2,7-DCF <10 0.1-5   0.5-2   3,5-DCF <10 <5 <2 9F-2CA <10 0.1-5   0.5-2  9F-4CA <5 <3 <1 Total fluorenones <40 <20 1-8 4,4′-DCB <45 <15 0.5-5  2,5,4′-TCB <45 0.1-15  0.5-5   PA <1,000  15-400  40-150 IPA 2,500  40-1,200 120-400 BA <4,500   50-1,500 150-500 TMA <1,000  15-400 40-150 2,6-DCBC <40 <20 <5 4,4′-DCBZ <40 <20 <5 4,4′-DCBP <40 <20 <52,5,4′-TCBP <40 <20 0.5-5   TPA <9,000   200-6,000   400-2,000Precipitated <9,000   200-6,000   600-2,000 Solids at 20° C. TotalAromatic <18,000   300-9,000   450-3,000 Carboxylic Acids Lacking Non-Aromatic Hydrocarbyl Groups

Many other aromatic impurities are also typically present in recycledsolvent, generally varying at even lower levels and/or in proportion toone or more of the disclosed aromatic compounds. Methods for controllingthe disclosed aromatic compounds in the preferred ranges will typicallykeep other aromatic impurities at suitable levels.

When bromine is used within the reaction medium, a large number of ionicand organic forms of bromine are known to exist in a dynamicequilibrium. These various forms of bromine have different stabilitycharacteristics once leaving the reaction medium and passing throughvarious unit operations pertaining to recycled solvent. For example,alpha-bromo-para-toluic acid may persist as such at some conditions ormay rapidly hydrolyze at other conditions to form 4-hydroxymethylbenzoicacid and hydrogen bromide. In the present invention, it is preferablethat at least about 40 weight percent, more preferable that at leastabout 60 weight percent, and most preferable that at least about 80weight percent of the total mass of bromine present in the aggregatedsolvent feed to the reaction medium is in one or more of the followingchemical forms: ionic bromine, alpha-bromo-para-toluic acid, andbromoacetic acid.

Although the importance and value of controlling the aggregated weightaverage purity of solvent feed within the disclosed, desired ranges ofthe present invention has not heretofore been discovered and/ordisclosed, suitable means for controlling the solvent feed purity may beassembled from various methods already known in the art. First, anysolvent evaporated from the reaction medium is typically of suitablepurity providing that liquid or solids from the reaction medium are notentrained with the evaporated solvent. The feeding of reflux solventdroplets into the off-gas disengaging space above the reaction medium,as disclosed herein, appropriately limits such entrainment; and recycledsolvent of suitable purity with respect to aromatic compound can becondensed from such off-gas. Second, the more difficult and costlypurification of recycled solvent feed typically relates to solvent takenfrom the reaction medium in liquid form and to solvent that subsequentlycontacts the liquid and/or solid phases of the reaction medium withdrawnfrom the reaction vessel (e.g., recycled solvent obtained from a filterin which solids are concentrated and/or washed, recycled solventobtained from a centrifuge in which solids are concentrated and/orwashed, recycled solvent taken from a crystallization operation, and soon). However, means are also known in the art for effecting thenecessary purification of these recycled solvent streams using one ormore prior disclosures. With respect to controlling precipitated solidsin recycled solvent to be within the ranges specified, suitable controlmeans include, but are not limited to, gravimetric sedimentation,mechanical filtration using filter cloth on rotary belt filters androtary drum filters, mechanical filtration using stationary filtermedium within pressure vessels, hydro-cyclones, and centrifuges. Withrespect to controlling dissolved aromatic species in recycled solvent tobe within the ranges specified, the control means include, but are notlimited to, those disclosed in U.S. Pat. No. 4,939,297 and U.S. patentapplication Pub. No. 2005-0038288, incorporated herein by reference.However, none of these prior inventions discovered and disclosed thepreferred levels of purity in the aggregated solvent feed as disclosedherein. Rather, these prior inventions merely provided means to purifyselected and partial streams of recycled solvent without deducing thepresent inventive, optimal values of the composition of the aggregatedweight average solvent feed to the reaction medium.

Turning now to the purity of the feed of oxidizable compound, it isknown that certain levels of isophthalic acid, phthalic acid, andbenzoic acid are present and tolerable at low levels in purified TPAused for polymer production. Moreover, it is known these species arerelatively more soluble in many solvents and may be advantageouslyremoved from purified TPA by crystallization processes. However, from anembodiment of the invention disclosed herein, it is now known thatcontrolling the level of several relatively soluble aromatic species,notably including isophthalic acid, phthalic acid, and benzoic acid, inthe liquid phase of the reaction medium is surprisingly important forcontrolling the level of polycyclic and colored aromatic compoundscreated in the reaction medium, for controlling compounds with more than2 carboxylic acid functions per molecule, for controlling reactionactivity within the partial oxidation reaction medium, and forcontrolling yield losses of oxidant and of aromatic compound.

It is known within the art that isophthalic acid, phthalic acid, andbenzoic acid are formed in the reaction medium as follows. Meta-Xylenefeed impurity oxidizes in good conversion and yield to IPA. Ortho-Xylenefeed impurity oxidizes in good conversion and yield to phthalic acid.Ethylbenzene and toluene feed impurities oxidize in good conversion andyield to benzoic acid. However, the inventors have observed thatsignificant amounts of isophthalic acid, phthalic acid, and benzoic acidare also formed within a reaction medium comprising para-xylene by meansother than oxidation of meta-xylene, ortho-xylene, ethylbenzene, andtoluene. These other intrinsic chemical routes possibly includedecarbonylation, decarboxylation, the re-organization of transitionstates, and addition of methyl and carbonyl radicals to aromatic rings.

In determining preferred ranges of impurities in the feed of oxidizablecompound, many factors are relevant. Any impurity in the feed is likelyto be a direct yield loss and a product purification cost if the purityrequirements of the oxidized product are sufficiently strict (e.g., in areaction medium for partial oxidation of para-xylene, toluene andethylbenzene typically found in commercial-purity para-xylene lead tobenzoic acid, and this benzoic acid is largely removed from mostcommercial TPA). When the partial oxidation product of a feed impurityparticipates in additional reactions, factors other than simple yieldloss and removal become appropriate when considering how much feedpurification cost to incur (e.g., in a reaction medium for partialoxidation of para-xylene, ethylbenzene leads to benzoic acid, andbenzoic acid subsequently leads to highly colored9-fluorenone-2-carboxylic acid, to isophthalic acid, to phthalic acid,and to increased carbon oxides, among others). When the reaction mediumself-generates additional amounts of an impurity by chemical mechanismsnot directly related to feed impurities, the analysis becomes still morecomplex (e.g., in a reaction medium for partial oxidation ofpara-xylene, benzoic acid is also self-generated from para-xyleneitself). In addition, the downstream processing of the crude oxidationproduct may affect the considerations for preferred feed purity. Forexample, the cost of removing to suitable levels a direct impurity(benzoic acid) and subsequent impurities (isophthalic acid, phthalicacid, 9-fluorenone-2-carboxylic acid, et al.) may be one and the same,may be different from each other, and may be different from therequirements of removing a largely unrelated impurity (e.g., incompleteoxidation product 4-CBA in the oxidation of para-xylene to TPA).

The following disclosed feed purity ranges for para-xylene are preferredwhere para-xylene is fed with solvent and oxidant to a reaction mediumfor partial oxidation to produce TPA. These ranges are more preferred inTPA production process having post-oxidation steps to remove fromreaction medium impurities other than oxidant and solvent (e.g.,catalyst metals). These ranges are still more preferred in TPAproduction processes that remove additional 4-CBA from CTA (e.g., byconversion of CTA to dimethyl terephthalate plus impurity esters andsubsequent separation of the methyl ester of 4-CBA by distillation, byoxidative digestion methods for converting 4-CBA to TPA, byhydrogenation methods for converting 4-CBA to para-toluic acid, which isthen separated by partial-crystallization methods). These ranges aremost preferred in TPA production processes that remove additional 4-CBAfrom CTA by oxidative digestion methods for converting 4-CBA to TPA.

Using new knowledge of preferred ranges of recycling aromatic compoundsand of the relative amounts of the aromatic compounds formed directlyfrom oxidation of feed impurities as compared to other intrinsicchemical routes, improved ranges for impurities have been discovered forimpure para-xylene being fed to a partial oxidation process for TPAproduction. Table 2, below provides preferred values for the amount ofmeta-xylene, ortho-xylene, and ethylbenzene+toluene in the para-xylenefeed. TABLE 2 Components of Impure para-xylene Feed Preferred ComponentAmt. More Preferred Most Preferred Identification (ppmw) Amt. (ppmw)Amt. (ppmw) meta-xylene 20-800 50-600 100-400 ortho-xylene 10-300 20-200 30-100 ethylbenzene + toluene* 20-700 50-500 100-300 total 50-900100-800  200-700*Specification for ethylbenzene + toluene is each separately and in sum

Those skilled in the art will now recognize the above impurities withinimpure para-xylene may have their greatest effect on the reaction mediumafter their partial oxidation products have accumulated in recycledsolvent. For example, feeding the upper amount of the most preferredrange of meta-xylene, 400 ppmw, will immediately produce about 200 ppmwof isophthalic acid within the liquid phase of the reaction medium whenoperating with about 33 weight percent solids in the reaction medium.This compares with an input from the upper amount of the most preferredrange for isophthalic acid in recycled solvent of 400 ppmw which, afterallowing for a typical solvent evaporation to cool the reaction medium,amounts to about 1,200 ppmw of isophthalic acid within the liquid phaseof the reaction medium. Thus, it is the accumulation of partialoxidation products over time within recycled solvent that represents thegreatest probable impact of the meta-xylene, ortho-xylene, ethylbenzene,and toluene impurities in the feed of impure para-xylene. Accordingly,the above ranges for impurities in impure para-xylene feed are preferredto be maintained for at least one-half of each day of operation of anypartial oxidation reaction medium in a particular manufacturing unit,more preferably for at least three-quarters of each day for at leastseven consecutive days of operation, and most preferably when themass-weighted averages of the impure para-xylene feed composition arewithin the preferred ranges for at least 30 consecutive days ofoperation.

Means for obtaining impure para-xylene of preferred purity are alreadyknown in the art and include, but are not limited to, distillation,partial crystallization methods at sub-ambient temperatures, andmolecular sieve methods using selective pore-size adsorption. However,the preferred ranges of purity specified herein are, at their high end,more demanding and expensive than characteristically practiced bycommercial suppliers of para-xylene; and yet at the low end, thepreferred ranges avoid overly costly purification of para-xylene forfeeding to a partial oxidation reaction medium by discovering anddisclosing where the combined effects of impurity self-generation frompara-xylene itself and of impurity consumptive reactions within thereaction medium become more important than the feed rates of impuritieswithin impure para-xylene.

When the xylene-containing feed stream contains selected impurities,such as ethyl-benzene and/or toluene, oxidation of these impurities cangenerate benzoic acid. As used herein, the term “impurity-generatedbenzoic acid” shall denote benzoic acid derived from any source otherthan xylene during xylene oxidation.

As disclosed herein, a portion of the benzoic acid produced duringxylene oxidation is derived from the xylene itself. This production ofbenzoic acid from xylene is distinctly in addition to any portion ofbenzoic acid production that may be impurity-generated benzoic acid.Without being bound by theory, it is believed that benzoic acid isderived from xylene within the reaction medium when various intermediateoxidation products of xylene spontaneously decarbonylate (carbonmonoxide loss) or decarboxylate (carbon dioxide loss) to thereby producearyl radicals. These aryl radicals can then abstract a hydrogen atomfrom one of many available sources in the reaction medium and produceself-generated benzoic acid. Whatever the chemical mechanism, the term“self-generated benzoic acid”, as used herein, shall denote benzoic acidderived from xylene during xylene oxidation.

As also disclosed herein, when para-xylene is oxidized to produceterephthalic acid (TPA), the production of self-generated benzoic acidcauses para-xylene yield loss and oxidant yield loss. In addition, thepresence of self-generated benzoic acid in the liquid phase of thereaction medium correlates with increases for many undesirable sidereactions, notably including generation of highly colored compoundscalled mono-carboxy-fluorenones. Self-generated benzoic acid alsocontributes to the undesirable accumulation of benzoic acid in recycledfiltrate which further elevates the concentration of benzoic acid in theliquid phase of the reaction medium. Thus, formation of self-generatedbenzoic acid is desirably minimized, but this is also appropriatelyconsidered simultaneously with impurity-generated benzoic acid, withfactors affecting consumption of benzoic acid, with factors pertainingto other issues of reaction selectivity, and with overall economics.

The inventors have discovered that the self-generation of benzoic acidcan be controlled to low levels by appropriate selection of, forexample, temperature, xylene distribution, and oxygen availabilitywithin the reaction medium during oxidation. Not wishing to be bound bytheory, lower temperatures and improved oxygen availability appear tosuppress the decarbonylation and/or decarboxylation rates, thus avoidingthe yield loss aspect of self-generated benzoic acid. Sufficient oxygenavailability appears to direct aryl radicals toward other more benignproducts, in particular hydroxybenzoic acids. Distribution of xylene inthe reaction medium may also affect the balance between aryl radicalconversion to benzoic acid or to hydroxybenzoic acids. Whatever thechemical mechanisms, the inventors have discovered reaction conditionsthat, although mild enough to reduce benzoic acid production, are severeenough to oxidize a high fraction of the hydroxybenzoic acid productionto carbon monoxide and/or carbon dioxide, which are easily removed fromthe oxidation product.

In a preferred embodiment of the present invention, the oxidationreactor is configured and operated in a manner such that the formationof self-generated benzoic acid is minimized and the oxidation ofhydroxybenzoic acids to carbon monoxide and/or carbon dioxide ismaximized. When the oxidation reactor is employed to oxidize para-xyleneto terephthalic acid, it is preferred that para-xylene makes up at leastabout 50 weight percent of the total xylene in the feed streamintroduced into the reactor. More preferably, para-xylene makes up atleast about 75 weight percent of the total xylene in the feed stream.Still more preferably, para-xylene makes up at least 95 weight percentof the total xylene in the feed stream. Most preferably, para-xylenemakes up substantially all of the total xylene in the feed stream.

When the reactor is employed to oxidize para-xylene to terephthalicacid, it is preferred for the rate of production of terephthalic acid tobe maximized, while the rate of production of self-generated benzoicacid is minimized. Preferably, the ratio of the rate of production (byweight) of terephthalic acid to the rate of production (by weight) ofself-generated benzoic acid is at least about 500:1, more preferably atleast about 1,000:1, and most preferably at least 1,500:1. As will beseen below, the rate of production of self-generated benzoic acid ispreferably measured when the concentration of benzoic acid in the liquidphase of the reaction medium is below 2,000 ppmw, more preferably below1,000 ppmw, and most preferably below 500 ppmw, because these lowconcentrations suppress to suitably low rates reactions that convertbenzoic acid to other compounds.

Combining the self-generated benzoic acid and the impurity-generatedbenzoic acid, the ratio of the rate of production (by weight) ofterephthalic acid to the rate of production (by weight) of total benzoicacid is preferably at least about 400:1, more preferably at least about700:1, and most preferably at least 1,100:1. As will be seen below, thesummed rate of production of self-generated benzoic acid plusimpurity-generated benzoic acid is preferably measured when theconcentration of benzoic acid in the liquid phase of the reaction mediumis below 2,000 ppmw, more preferably below 1,000 ppmw, and mostpreferably below 500 ppmw, because these low concentrations suppress tosuitably low rates reactions that convert benzoic acid to othercompounds.

As disclosed herein, elevated concentrations of benzoic acid in theliquid phase of the reaction medium lead to increased formation of manyother aromatic compounds, several of which are noxious impurities inTPA; and, as disclosed herein, elevated concentrations of benzoic acidin the liquid phase of the reaction medium lead to increased formationof carbon oxide gases, the formation of which represents yield loss onoxidant and on aromatic compounds and/or solvent. Furthermore, it is nowdisclosed that the inventors have discovered a considerable portion ofthis increased formation of other aromatic compounds and of carbonoxides derives from reactions that convert some of the benzoic acidmolecules themselves, as contrasted to benzoic acid catalyzing otherreactions without itself being consumed. Accordingly, the “netgeneration of benzoic acid” is defined herein as the time-averagedweight of all benzoic acid exiting the reaction medium minus thetime-averaged weight of all benzoic acid entering the reaction mediumduring the same period of time. This net generation of benzoic acid isoften positive, driven by the formation rates of impurity-generatedbenzoic acid and of self-generated benzoic acid. However, the inventorshave discovered that the conversion rate of benzoic acid to carbonoxides, and to several other compounds, appears to increaseapproximately linearly as the concentration of benzoic acid is increasedin the liquid phase of the reaction medium, measured when other reactionconditions comprising temperature, oxygen availability, STR, andreaction activity are maintained appropriately constant. Thus, when theconcentration of benzoic acid in the liquid-phase of the reaction mediumis great enough, perhaps due to an elevated concentration of benzoicacid in recycled solvent, then the conversion of benzoic acid moleculesto other compounds, including carbon oxides, can become equal to orgreater than the chemical generation of new benzoic acid molecules. Inthis case, the net generation of benzoic acid can become balanced nearzero or even negative. The inventors have discovered that when the netgeneration of benzoic acid is positive, then the ratio of the rate ofproduction (by weight) of terephthalic acid in the reaction mediumcompared to the rate of net generation of benzoic acid in the reactionmedium is preferably above about 700:1, more preferably above about1,100:1, and most preferably above 4,000:1. The inventors havediscovered that when the net generation of benzoic acid is negative, theratio of the rate of production (by weight) of terephthalic acid in thereaction medium compared to the rate of net generation of benzoic acidin the reaction medium is preferably above about 200:(−1), morepreferably above about 1,000:(−1), and most preferably above 5,000:(−1).

The inventors have also discovered preferred ranges for the compositionof the slurry (liquid+solid) withdrawn from the reaction medium and forthe solid CTA portion of the slurry. The preferred slurry and thepreferred CTA compositions are surprisingly superior and useful. Forexample, purified TPA produced from this preferred CTA by oxidativedigestion has a sufficiently low level of total impurities and ofcolored impurities such that the purified TPA is suitable, withouthydrogenation of additional 4-CBA and/or colored impurities, for a widerange of applications in PET fibers and PET packaging applications. Forexample, the preferred slurry composition provides a liquid phase of thereaction medium that is relatively low in concentration of importantimpurities and this importantly reduces the creation of other even moreundesirable impurities as disclosed herein. In addition, the preferredslurry composition importantly aids the subsequent processing of liquidfrom the slurry to become suitably pure recycled solvent, according toother embodiments of the present invention.

CTA produced according to one embodiment of the present inventioncontains less impurities of selected types than CTA produce byconventional processes and apparatuses, notably those employing recycledsolvent. Impurities that may be present in CTA include the following:4-carboxybenzaldehyde (4-CBA), 4,4′-dicarboxystilbene (4,4′-DCS),2,6-dicarboxyanthraquinone (2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF),2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5DCF),9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid(9F-4CA), 4,4′-dicarboxybiphenyl (4,4′-DCB), 2,5,4′-tricarboxybiphenyl(2,5,4′-TCB), phthalic acid (PA), isophthalic acid (IPA), benzoic acid(BA), trimellitic acid (TMA), para-toluic acid (PTAC),2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4′-dicarboxybenzil (4,4′-DCBZ),4,4′-dicarboxybenzophenone (4,4′-DCBP), 2,5,4′-tricarboxybenzophenone(2,5,4′-TCBP). Table 3, below provides the preferred amounts of theseimpurities in CTA produced according to an embodiment of the presentinvention. TABLE 3 CTA Impurities Impurity Preferred More Preferred MostPreferred Identification Amt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA<15,000   100-8,000  400-2,000 4,4′-DCS <12 <6 <3 2,6-DCA <9 <6 <22,6-DCF <100  2-50 5-25 2,7-DCF <30 <15 <5 3,5-DCF <16 <8 <2 9F-2CA <16<8 <4 9F-4CA <8 <4 <2 Total fluorenones <100  2-60 4-35 4,4′-DCB <64 1-32 2-8  2,5,4′-TCB <24 <12 <8 PA <200  3-100 5-50 IPA <800  10-40020-200 BA <600  5-300 15-100 TMA <800  10-400 20-200 PTAC <2,000  10-1,000 50-500 2,6-DCBC <64 <32 <8 4,4′-DCBZ <12 <8 <4 4,4′-DCBP <40<30 <20 2,5,4′-TCBP <32 <16 <4

In addition, it is preferred for CTA produced according to an embodimentof the present invention to have reduced color content relative to CTAproduce by conventional processes and apparatuses, notably thoseemploying recycled solvent. Thus, it is preferred for CTA produced inaccordance to one embodiment of the present invention have a percenttransmittance percent at 340 nanometers (nm) of at least about 25percent, more preferably of at least about 50 percent, and mostpreferably of at least 60 percent. It is further preferred for CTAproduced in accordance to one embodiment of the present invention tohave a percent transmittance percent at 400 nanometers (nm) of at leastabout 88 percent, more preferably of at least about 90 percent, and mostpreferably of at least 92 percent.

The test for percent transmittance provides a measure of the colored,light-absorbing impurities present within TPA or CTA. As used herein,the test refers to measurements done on a portion of a solution preparedby dissolving 2.00 grams of dry solid TPA or CTA in 20.0 milliliters ofdimethyl sulfoxide (DMSO), analytical grade or better. A portion of thissolution is then placed in a Hellma semi-micro flow cell, PN 176.700,which is made of quartz and has a light path of 1.0 cm and a volume of0.39 milliliters. (Hellma USA, 80 Skyline Drive, Plainview, N.Y. 11803).An Agilent 8453 Diode Array Spectrophotometer is used to measure thetransmittance of different wavelengths of light through this filled flowcell. (Agilent Technologies, 395 Page Mill Road, Palo Alto, Calif.94303). After appropriate correction for absorbance from the background,including but not limited to the cell and the solvent used, the percenttransmittance results, characterizing the fraction of incident lightthat is transmitted through the solution, are reported directly by themachine. Percent transmittance values at light wavelengths of 340nanometers and 400 nanometers are particularly useful for discriminatingpure TPA from many of the impurities typically found therein.

The preferred ranges of various aromatic impurities in the slurry(solid+liquid) phase of the reaction medium are provided below in Table4. TABLE 4 Slurry Impurities Impurity Preferred More Preferred MostPreferred Identification Amt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA<8,000 <5,000 <2,500 4,4′-DCS <4 <2 <1 2,6-DCA <6 <3 <1 2,6-DCF <70 2-404-20 2,7-DCF <12 <8 <4 3,5-DCF <12 <8 <4 9F-2CA <12 <8 <4 9F-4CA <8 <4<2 Total fluorenones <90 2-60 5-30 4,4′-DCB <64 1-16 2-4  2,5,4′-TCB <602-40 4-20 PA <3,000   25-1,500 75-500 IPA 9,000   75-4,500  225-1,500 BA<15,000  100-6,000  300-2,000 TMA <3,000   25-1,500 75-500 PTAC <8,000 100-4,000  200-2,000 4,4′-DCBZ <5 <4 <3 4,4′-DCBP <240 <160 <802,5,4′-TCBP <120 <80 <40

These preferred compositions for the slurry embody the preferredcomposition of the liquid phase of the reaction medium while usefullyavoiding experimental difficulties pertaining to precipitation ofadditional liquid phase components from the reaction medium into solidphase components during sampling from the reaction medium, separation ofliquids and solids, and shifting to analytical conditions.

Many other aromatic impurities are also typically present in the slurryphase of the reaction medium and in CTA of the reaction medium,generally varying at even lower levels and/or in proportion to one ormore of the disclosed aromatic compounds. Controlling the disclosedaromatic compounds in the preferred ranges will keep other aromaticimpurities at suitable levels. These advantaged compositions for theslurry phase in the reaction medium and for the solid CTA taken directlyfrom the slurry are enabled by operating with embodiments of theinvention disclosed herein for partial oxidation of para-xylene to TPA.

Measurement of the concentration of low level components in the solvent,recycled solvent, CTA, slurry from the reaction medium, and PTA areperformed using liquid chromatography methods. Two interchangeableembodiments are now described.

The method referred to herein as HPLC-DAD comprises high pressure liquidchromatography (HPLC) coupled with a diode array detector (DAD) toprovide separation and quantitation of various molecular species withina given sample. The instrument used in this measurement is a model 1100HPLC equipped with a DAD, provided by Agilent Technologies (Palo Alto,Calif.), though other suitable instruments are also commerciallyavailable and from other suppliers As is known in the art, both theelution time and the detector response are calibrated using knowncompounds present in known amounts, compounds and amounts that areappropriate to those occurring in actual unknown samples.

The method referred to herein as HPLC-MS comprises high pressure liquidchromatography (HPLC) coupled with mass spectrometry (MS) to provideseparation, identification, and quantitation of various molecularspecies within a given sample. The instruments used in this measurementis an Alliance HPLC and ZQ MS provided by Waters Corp. (Milford, Mass.),though other suitable instruments are also commercially available andfrom other suppliers. As is known in the art, both the elution time andthe mass spectrometric response are calibrated using known compoundspresent in known amounts, compounds and amounts that are appropriate tothose occurring in actual unknown samples.

Another embodiment of the current invention relates to partial oxidationof aromatic oxidizable compound with appropriate balancing of thesuppression of noxious aromatic impurities on the one hand against theproduction of carbon dioxide and carbon monoxide, collectively carbonoxides (COx), on the other. These carbon oxides typically exit thereaction vessel in the off-gas, and they correspond to a destructiveloss of solvent and of oxidizable compound, including the ultimatelypreferred oxidized derivatives (e.g., acetic acid, para-xylene, andTPA). The inventors have discovered lower bounds for the production ofcarbon oxides below which it seems the high creation of noxious aromaticimpurities, as described below, and the low overall conversion level areinevitably too poor to be of economic utility. The inventors have alsodiscovered upper bounds of carbon oxides above which the generation ofcarbon oxides continues to increase with little further value providedby reduction in generation of noxious aromatic impurities.

The inventors have discovered that reducing the liquid-phaseconcentrations of aromatic oxidizable compound feed and of aromaticintermediate species within a reaction medium leads to lower generationrates for noxious impurities during the partial oxidation of aromaticoxidizable compound. These noxious impurities include coupled aromaticrings and/or aromatic molecules containing more than the desired numberof carboxylic acid groups (e.g., in the oxidation of para-xylene thenoxious impurities include 2,6-dicarboxyanthraquinone,2,6-dicarboxyfluorenone, trimellitic acid, 2,5,4′-tricarboxybiphenyl,and 2,5,4′-benzophenone). The aromatic intermediate species includearomatic compounds descended from the feed of oxidizable aromaticcompound and still retaining non-aromatic hydrocarbyl groups (e.g., inthe oxidation of para-xylene the aromatic intermediate species comprisepara-tolualdehyde, terephthaldehyde, para-toluic acid, 4-CBA,4-hydroxymethylbenzoic acid, and alpha-bromo-para-toluic acid). Thearomatic oxidizable compound feed and the aromatic intermediate speciesretaining non-aromatic hydrocarbyl groups, when present in the liquidphase of the reaction medium, appear to lead to noxious impurities in amanner similar to that already disclosed herein for dissolved aromaticspecies lacking non-aromatic hydrocarbyl groups (e.g., isophthalicacid).

Set against this need for higher reaction activity to suppress formationof noxious aromatic impurities during partial oxidation of oxidizablearomatic compound, the inventors have discovered that the undesirableattendant result is increased production of carbon oxides. It isimportant to appreciate that these carbon oxides represent a yield lossof oxidizable compound and oxidant, not just solvent. Explicitly, asubstantial and sometimes principal fraction of the carbon oxides comesfrom the oxidizable compound, and its derivatives, rather than fromsolvent; and often the oxidizable compound costs more per carbon unitthan does solvent. Furthermore, it is important to appreciate that thedesired product carboxylic acid (e.g., TPA) is also subject toover-oxidation to carbon oxides when present in the liquid phase of thereaction medium.

It is also important to appreciate that the present invention relates toreactions in the liquid phase of the reaction medium and to reactantconcentrations therein. This is in contrast to some prior inventionswhich relate directly to the creation in precipitated solid form ofaromatic compound retaining non-aromatic hydrocarbyl groups.Specifically, for the partial oxidation of para-xylene to TPA, certainprior inventions pertain to the amount of 4-CBA precipitated in thesolid phase of CTA. However, the present inventors have discovered avariance of greater than two to one for the ratio of 4-CBA in the solidphase to 4-CBA in the liquid phase, using the same specifications oftemperature, pressure, catalysis, solvent composition and space-timereaction rate of para-xylene, depending upon whether the partialoxidation is conducted in a well-mixed autoclave or in a reaction mediumwith oxygen and para-xylene staging according to the present invention.Further, the inventors have observed that the ratio of 4-CBA in thesolid phase to 4-CBA in the liquid phase can also vary by over two toone in either well-mixed or staged reaction medium depending upon thespace-time reaction rate of para-xylene at otherwise similarspecifications of temperature, pressure, catalysis, and solventcomposition. Additionally, 4-CBA in the solid phase CTA does not appearto contribute to the formation of noxious impurities, and 4-CBA in thesolid phase can be recovered and oxidized on to TPA simply and at highyield (e.g., by oxidative digestion of the CTA slurry as is describedherein); whereas the removal of noxious impurities is far more difficultand costly than removal of solid phase 4-CBA, and the production ofcarbon oxides represents a permanent yield loss. Thus, it is importantto distinguish that this aspect of the present invention relates toliquid-phase compositions in the reaction medium.

Whether sourced from solvent or oxidizable compound, the inventors havediscovered that at conversions of commercial utility the production ofcarbon oxides relates strongly to the level of overall reaction activitydespite wide variation in the specific combination of temperature,metals, halogens, temperature, acidity of the reaction medium asmeasured by pH, water concentration employed to obtain the level ofoverall reaction activity. The inventors have found it useful for thepartial oxidation of xylene to evaluate the level of overall reactionactivity using the liquid-phase concentration of toluic acids at themid-height of the reaction medium, the bottom of the reaction medium,and the top of the reaction medium.

Thus, there arises an important simultaneous balancing to minimize thecreation of noxious impurities by increasing reaction activity and yetto minimize the creation of carbon oxides by lowering reaction activity.That is, if the overall production of carbon oxides is suppressed toolow, then excessive levels of noxious impurities are formed, and viceversa.

Furthermore, the inventors have discovered that the solubility and therelative reactivity of the desired carboxylic acid (e.g., TPA) and thepresence of other dissolved aromatic species lacking non-aromatichydrocarbyl groups introduce a very important fulcrum in this balancingof carbon oxides versus noxious impurities. The desired productcarboxylic acid is typically dissolved in the liquid phase of thereaction medium, even when also present in solid form. For example, attemperatures in the preferred ranges, TPA is soluble in a reactionmedium comprising acetic acid and water at levels ranging from about onethousand ppmw to in excess of 1 weight percent, with solubilityincreasing as temperature increases. Notwithstanding that there aredifferences in the reaction rates toward forming various noxiousimpurities from oxidizable aromatic compound feed (e.g., para-xylene),from aromatic reaction intermediates (e.g., para-toluic acid), from thedesired product aromatic carboxylic acid (e.g., TPA), and from aromaticspecies lacking non-aromatic hydrocarbyl groups (e.g., isophthalicacid), the presence and reactivity of the latter two groups establishesa region of diminishing returns with regards to further suppression ofthe former two groups, oxidizable aromatic compound feed and aromaticreaction intermediates. For example, in a partial oxidation ofpara-xylene to TPA, if dissolved TPA amounts to 7,000 ppmw in the liquidphase of the reaction medium at given conditions, dissolved benzoic acidamounts to 8,000 ppmw, dissolved isophthalic acid amounts to 6,000 ppmw,and dissolved phthalic acid amounts to 2,000 ppmw, then the value towardfurther lowering of total noxious compounds begins to diminish asreaction activity is increased to suppress the liquid-phaseconcentration para-toluic acid and 4-CBA below similar levels. That is,the presence and concentration in the liquid phase of the reactionmedium of aromatic species lacking non-aromatic hydrocarbyl groups isvery little altered by increasing reaction activity, and their presenceserves to expand upwards the region of diminishing returns for reducingthe concentration of reaction intermediates in order to suppressformation of noxious impurities.

Thus, one embodiment of the present invention provides preferred rangesof carbon oxides, bounded on the lower end by low reaction activity andexcessive formation of noxious impurities and on upper end by excessivecarbon losses, but at levels lower than previously discovered anddisclosed as commercially useful. Accordingly, the formation of carbonoxides is preferably controlled as follows. The ratio of moles of totalcarbon oxides produced to moles of oxidizable aromatic compound fed ispreferably greater than about 0.02:1, more preferably greater than about0.04:1, still more preferably greater than about 0.05:1, and mostpreferably greater than 0.06:1. At the same time, the ratio of moles oftotal carbon oxides produced to moles of oxidizable aromatic compoundfed is preferably less than about 0.24:1, more preferably less thanabout 0.22:1, still more preferably less than about 0.19:1, and mostpreferably less than 0.15:1. The ratio of moles of carbon dioxideproduced to moles of oxidizable aromatic compound fed is preferablygreater than about 0.01:1, more preferably greater than about 0.03:1,still more preferably greater than about 0.04:1, and most preferablygreater than 0.05:1. At the same time, the ratio of moles of carbondioxide produced to moles of oxidizable aromatic compound fed ispreferably less than about 0.21:1, more preferably less than about0.19:1, still more preferably less than about 0.16:1, and mostpreferably less than 0.11. The ratio of moles of carbon monoxideproduced to moles of oxidizable aromatic compound fed is preferablygreater than about 0.005:1, more preferably greater than about 0.010:1,still more preferably greater than about 0.015:1, and most preferablygreater than 0.020:1. At the same time, the ratio of moles of carbonmonoxide produced to moles of oxidizable aromatic compound fed ispreferably less than about 0.09:1, more preferably less than about0.07:1, still more preferably less than about 0.05:1, and mostpreferably less than 0.04:1

The content of carbon dioxide in dry off-gas from the oxidation reactoris preferably greater than about 0.10 mole percent, more preferablygreater than about 0.20 mole percent, still more preferably greater thanabout 0.25 mole percent, and most preferably greater than 0.30 molepercent. At the same time, the content of carbon dioxide in dry off-gasfrom the oxidation reactor is preferably less than about 1.5 molepercent, more preferably less than about 1.2 mole percent, still morepreferably less than about 0.9 mole percent, and most preferably lessthan 0.8 mole percent. The content of carbon monoxide in dry off-gasfrom the oxidation reactor is preferably greater than about 0.05 molepercent, more preferably greater than about 0.10 mole percent, stillmore preferably greater than 0.15, and most preferably greater than 0.18mole percent. At the same time, the content of carbon monoxide in dryoff-gas from the oxidation reactor is preferably less than about 0.60mole percent, more preferably less than about 0.50 mole percent, stillmore preferably less than about 0.35 mole percent, and most preferablyless than 0.28 mole percent The inventors have discovered that animportant factor for reducing the production of carbon oxides to thesepreferred ranges is improving the purity of the recycled filtrate and ofthe feed of oxidizable compound to reduce the concentration of aromaticcompounds lacking non-aromatic hydrocarbyl groups according todisclosures of the present invention—this simultaneously reduces theformation of carbon oxides and of noxious impurities. Another factor isimproving distribution of para-xylene and oxidant within the reactionvessel according to disclosures of the present invention. Other factorsenabling the above preferred levels of carbon oxides are to operate withthe gradients in the reaction medium as disclosed herein for pressure,for temperature, for concentration of oxidizable compound in the liquidphase, and for oxidant in the gas phase. Other factors enabling theabove preferred levels of carbon oxides are to operate within thedisclosures herein preferred for space-time reaction rate, pressure,temperature, solvent composition, catalyst composition, and mechanicalgeometry of the reaction vessel.

An important benefit from operating within the preferred ranges ofcarbon oxide formation is that the usage of molecular oxygen can bereduced, though not to stoichiometric values. Notwithstanding the goodstaging of oxidant and oxidizable compound according to the presentinvention, an excess of oxygen must be retained above the stoichiometricvalue, as calculated for feed of oxidizable compound alone, to allow forsome losses to carbon oxides and to provide excess molecular oxygen tocontrol the formation of noxious impurities. Specifically for the casewhere xylene is the feed of oxidizable compound, the feed ratio ofweight of molecular oxygen to weight of xylene is preferably greaterthan about 0.91:1.00, more preferably greater than about 0.95:1.00, andmost preferably greater than 0.99:1.00. At the same time, the feed ratioof weight of molecular oxygen to weight of xylene is preferably lessthan about 1.20:1.00, more preferably less than about 1.12:1.00, andmost preferably less than 1.06:1.00. Specifically for xylene feed, thetime-averaged content of molecular oxygen in the dry off-gas from theoxidation reactor is preferably greater than about 0.1 mole percent,more preferably greater than about 1 mole percent, and most preferablygreater than 1.5 mole percent. At the same time, the time-averagedcontent of molecular oxygen in the dry off-gas from the oxidationreactor is preferably less than about 6 mole percent, more preferablyless than about 4 mole percent, and most preferably less than 3 molepercent.

Another important benefit from operating within the preferred ranges ofcarbon oxide formation is that less aromatic compound is converted tocarbon oxides and other less valuable forms. This benefit is evaluatedusing the sum of the moles of all aromatic compounds exiting thereaction medium divided by the sum of the moles of all aromaticcompounds entering the reaction medium over a continuous period of time,preferably one hour, more preferably one day, and most preferably 30consecutive days. This ratio is hereinafter referred to as the “molarsurvival ratio” for aromatic compounds through the reaction medium andis expressed as a numerical percentage. If all entering aromaticcompounds exit the reaction medium as aromatic compounds, albeit mostlyin oxidized forms of the entering aromatic compounds, then the molarsurvival ratio has its maximum value of 100 percent. If exactly 1 ofevery 100 entering aromatic molecules is converted to carbon oxidesand/or other non-aromatic molecules (e.g., acetic acid) while passingthrough reaction medium, then the molar survival ratio is 99 percent.Specifically for the case where xylene is the principal feed ofoxidizable aromatic compound, the molar survival ratio for aromaticcompounds through the reaction medium is preferably greater than about98 percent, more preferably greater than about 98.5 percent, and mostpreferably less than 99.0 percent. At the same time and in order thatsufficient overall reaction activity is present, the molar survivalratio for aromatic compounds through the reaction medium is preferablyless than about 99.9 percent, more preferably less than about 99.8percent, and most preferably less than 99.7 percent when xylene is theprincipal feed of oxidizable aromatic compound.

Another aspect of the current invention involves the production ofmethyl acetate in a reaction medium comprising acetic acid and one ormore oxidizable aromatic compounds. This methyl acetate is relativelyvolatile compared to water and acetic acid and thus tends to follow theoff-gas unless additional cooling or other unit operations are employedto recover it and/or to destroy it prior to releasing the off-gas backto the environment. The formation of methyl acetate thus represents anoperating cost and also a capital cost. Perhaps the methyl acetate isformed by first combining a methyl radical, perhaps from decompositionof acetic acid, with oxygen to produce methyl hydroperoxide, bysubsequently decomposing to form methanol, and by finally reacting theproduced methanol with remaining acetic acid to form methyl acetate.Whatever the chemical path, the inventors have discovered that whenevermethyl acetate production is at too low a rate, then the production ofcarbon oxides are also too low and the production of noxious aromaticimpurities are too high. If methyl acetate production is at too high arate, then the production of carbon oxides are also unnecessarily highleading to yield losses of solvent, oxidizable compound and oxidant.When employing the preferred embodiments disclosed herein, theproduction ratio of moles of methyl acetate produced to moles ofoxidizable aromatic compound fed is preferably greater than about0.005:1, more preferably greater than about 0.010:1, and most preferablygreater than 0.020:1. At the same time, the production ratio of moles ofmethyl acetate produced to moles of oxidizable aromatic compound fed ispreferably less than about 0.09:1, more preferably less than about0.07:1, still more preferably less than about 0.05:1, and mostpreferably less than 0.04:1.

Certain embodiments of this invention can be further illustrated by thefollowing examples, although it should be understood that these examplesare included merely for purposes of illustration and are not intended tolimit the scope of the invention unless otherwise specificallyindicated.

EXAMPLES 1-10

Example 1 is a calculated example of a bubble column oxidation reactorfor oxidizing para-xylene in a liquid phase of a three-phase reactionmedium. The bubble column reactor of Example 1 represents a provenindustrial design with a para-xylene feed rate of 7,000 kilograms perhour. Examples 2 through 10 are calculated examples for bubble columnoxidation reactors having operating capacities 7 times greater than thereactor of Example 1. FIG. 58 provides a table outlining the differentparameters of the bubble column oxidation reactor that are varied inExamples 1-10.

Example 1

This example employs a bubble column reaction vessel having a vertical,cylindrical section with an inside diameter equaling 2.44 meters. Theheight of the cylindrical section is 32 meters from the lower tangentline (TL) to the upper TL of the cylindrical section. The vessel isfitted with 2:1 elliptical heads at the top and bottom of thecylindrical section. The height from the bottom of the reaction mediumto the top of the cylindrical section is about 32.6 meters, and theoverall height of the reaction vessel is about 33.2 meters. Theoperating level is about 25.6 meters above the bottom of the reactionmedium.

Para-xylene is fed to the reactor at a steady rate of 7,000 kilogramsper hour. A filtrate solvent comprising primarily acetic acid is fedintimately commingled with the para-xylene at a steady rate of 70,000kilograms per hour. The feed is distributed within the reaction vesselnear an elevation of 2 meters above the bottom of the reaction mediumusing a horizontal distributor assembly designed to provide asubstantially uniform release of the feed over the area of cross sectionlying inside of a radius of 0.45*(inside diameter). The concentration ofcatalyst components in the filtrate solvent is such that the compositionwithin the liquid phase of the reaction medium is 1,800 ppmw of cobalt,1,800 ppmw of bromine, and 100 ppmw of manganese. A separate stream ofreflux solvent is fed as droplets into the gas-disengaging zone abovethe operating level of the reaction medium at a steady rate of 49,000kilograms per hour, and it is distributed over essentially all of thecross sectional area of the disengaging zone. This reflux solvent iswithout significant levels of catalyst components. The combined watercontent of the filtrate solvent feed and of the reflux solvent feed issuch that the concentration of water within the liquid phase of thereaction medium is 6 weight percent. The feed rate of air is steady at arate of 35,000 kilograms per hour through an oxidant inlet distributorsimilar to the one shown in FIGS. 12-15, and all of the oxidantadmission holes are located below the lower TL of the cylindricalsection. The operating pressure of the reaction vessel overhead gas issteadily 0.52 megapascal gauge. The reaction vessel is operated in asubstantially adiabatic manner so that the heat of reaction elevates thetemperature of the incoming feeds and evaporates much of the incomingsolvent. Measured near the mid-elevation of the reaction medium, theoperating temperature is about 160° C. Reaction medium comprising crudeterephthalic acid (CTA) is removed from the side of the reaction vesselat an elevation of 15 meters at a steady rate using an externalde-aeration vessel.

The L:D ratio is 13.4 and the H:W ratio is 10.5. The volume occupied bythe reaction medium is about 118 cubic meters, and the reaction vesselcontains about 58,000 kilograms of slurry. The ratio of slurry mass topara-xylene feed rate is about 8.3 hours. The space time reactionintensity is about 59 kilograms of para-xylene fed per cubic meter ofreaction medium per hour. The pressure differential from the bottom ofthe reaction medium to the overhead off-gas exiting the reaction vesselis about 0.12 megapascal. The superficial velocity of the gas phase atthe mid-height of the reaction medium is about 0.8 meters per second.The upright surface area in contact with the reaction medium is about197 square meters, which is about 3.16*Wmin*H. The ratio of the volumeof reaction medium to upright surface area in contact with the reactionmedium is about 0.60 meters. Under conditions of Example 1, it isestimated that decomposition of acetic acid in the reaction mediumamounts to approximately 0.03 kilograms per kilogram of produced CTA.This is a significant cost in overall production economics.

Useful indicators of end-to-end flow rates within the bubble columnreaction vessel are the maximum time-averaged upward velocities of theslurry and of the oxidant phase. In a cylindrical bubble column reactionvessel, these maximum upward velocities occur near the vertical axis ofsymmetry of the cylindrical section. Calculated by a method derived fromthe Gas-Liquid Recirculation Model of Gupta (reference Churn-turbulentbubble columns: experiments and modeling; Gupta, Puneet; WashingtonUniv., St. Louis, Mo., USA. Avail. UMI, Order No. DA3065044; (2002), 357pp. From: Diss. Abstr. Int., B 2003, 63(9), 4269. Dissertation writtenin English. CAN 140:130325 AN 2003:424187 CAPLUS) and using aproprietary data base, the maximum time-averaged upwards velocity of thegas phase near the mid-elevation of the reaction medium is about 3.1meters per second. Similarly calculated, the maximum upwards velocity,time-averaged, of the slurry near the mid-elevation of the reactionmedium is about 1.4 meters per second.

Another useful indicator of end-to-end flow rates within the bubblecolumn reaction vessel is the maximum time-averaged downward velocity ofthe slurry in the parts of the reaction vessel located away from thecentral core. In a cylindrical reaction vessel, this maximum downwardvelocity typically occurs in the region lying outside a radius of0.35*(inside diameter) from the vertical axis of symmetry of thecylindrical section. Calculated by a method derived from the Gas-LiquidRecirculation Model of Gupta and using a proprietary data base, themaximum time-averaged downwards velocity of the slurry in the outerannulus near the mid-elevation of the reaction medium is about 1.4meters per second.

Example 2

In this example, the bubble column reactor is fed para-xylene at anincreased rate of 49,000 kilograms per hour—7 times greater than inExample 1. The superficial gas velocity, often considered an importantscale-up variable for bubble columns, is kept approximately equal toExample 1 by increasing the cross-sectional area of the reaction vesselto be about 7 times larger than in Example 1. The H:W and L:D ratios,often considered important scale-up variables for bubble columns, arealso kept approximately equal to Example 1.

The other feed flows are increased with the same 7:1 ratio to Example 1.The compositions of the feeds are the same as in Example 1, providingthe same concentrations of water, cobalt, bromine and manganese withinthe liquid phase of the reaction medium as in Example 1. The operatingpressure of the reaction vessel overhead gas is again 0.52 megapascalgauge, and the operating temperature is again about 160° C. measurednear the mid-elevation of the reaction medium. Reaction mediumcomprising CTA is removed from the side of the reaction vessel at anelevation of 40 meters at a steady rate using an external de-aerationvessel.

The bubble column reaction vessel comprises a vertical, cylindricalsection with an inside diameter equal to 6.46 meters. The L:D ratio ismaintained the same as Example 1, and the height from the bottom of thereaction medium to the top of the cylindrical section is thus 86.3meters. The top and bottom of the cylindrical section are fitted with2:1 elliptical heads, and the overall height of the reaction vessel isvery tall, about 88.0 meters. The feed is again distributed within thereaction vessel near an elevation of 2 meters above the bottom of thereaction medium using a horizontal distributor assembly designed for asubstantially uniform release of the feed over the area of cross sectionlying inside of a radius of 0.45*(inside diameter). The feed of air isagain through an oxidant inlet distributor similar to the one shown inFIGS. 12-15, and all of the oxidant admission holes are located belowthe lower TL of the cylindrical section. The reflux solvent isdistributed as droplets over essentially all of the cross sectional areaof the disengaging zone.

The H:W ratio of the reaction medium is kept approximately equal toExample 1, and thus the operating level is about 67.8 meters of reactionmedium. This leaves a disengaging height of about 18.5 meters in thecylindrical section plus about 1.6 meters in the top elliptical head.This disengaging height is excessive by about 10 meters. Thus, scalingthe vessel with constant L:D produces a mechanical installation that isoverly expensive for capital (e.g., excessive costs for pressure vessel,for foundations due to mass and wind load, for structural steel, forprocess and utility piping, and/or for instrument and electricalcabling).

The volume occupied by the reaction medium is about 2,200 cubic meters,and the reaction vessel contains about 1,100,000 kilograms of slurry.The ratio of slurry mass to para-xylene feed rate is about 22 hours,greatly increased compared to Example 1. The space time reactionintensity is only about 22 kilograms of para-xylene fed per cubic meterof reaction medium per hour, greatly decreased compared to Example 1.The pressure differential from the bottom of the reaction medium to theoverhead off-gas exiting the reaction vessel rises to about 0.33megapascal, greatly increased compared to Example 1. The upright surfacearea in contact with the reaction medium is about 1,393 square meters,which is about 3.18*Wmin*H. The ratio of the volume of reaction mediumto upright surface area in contact with the reaction medium is about1.58 meters, greatly increased compared to Example 1.

Thus, scaling the reaction medium dimensions to maintain bothsuperficial velocity and H:W ratio as approximately constant betweenthis example and Example 1 has produced very large changes in thereaction conditions. On balance, these changes are highly unfavorable.There are positive effects in this example (e.g., more diluteconcentrations of oxidizable compound, lower demand on mass transferrate per unit volume for molecular oxygen from gas to liquid phase, andso on) leading to production of fewer undesirable, colored byproductsper unit CTA. However, there are severe economic penalties relating tothe decomposition of acetic acid and to the pressure and power requiredto supply air to the bottom of the bubble column reaction vessel. Thedecomposition of acetic acid is approximately proportional to the massof acetic acid in the reaction medium, whenever the operatingtemperature and the composition of the liquid phase are keptapproximately constant and when para-xylene is fed with an excess ofmolecular oxygen. Owing to the much greater mass of acetic acid in thereaction vessel compared to the amount of CTA produced in this example,it is estimated that decomposition of acetic acid rises to about 0.09kilograms per kilogram of produced CTA. In addition, the air compressormust deliver air into the reaction medium with a pressure that is 0.85megapascal gauge in this example, whereas the delivered pressure is 0.64megapascal gauge in Example 1. For an air delivery rate of 245,000kilograms per hour and with typical allowance for various compressionand delivery efficiencies, the additional power requirement for thehigher delivery pressure in this example is about 3,000 kilowatts,continuously. Thus, scaling the reaction medium of this example withapproximately constant superficial gas velocity and H:W ratio providesunacceptable economics, despite the good quality of CTA expected.

Example 3

This example scales the process of Example 1 using superficial velocityand space-time reaction intensity. This leads to poor product qualitybecause, in simple terms, the natural convection flow patternsinherently produce a poor reaction profile vertically.

In this example, the feed rate of para-xylene is again 49,000 kilogramsper hour—7 times larger than in Example 1. The superficial gas velocityis again kept approximately equal to Example 1, but the L:D and H:Wratios are not kept equal. Instead, the STR is kept approximately equalto Example 1. This provides a column base pressure and a decompositionratio of acetic acid that are approximately equal to Example 1. Theother feed flows are increased with the same 7:1 ratio to Example 1. Thecompositions of the feeds are the same as in Example 1, providing thesame concentrations of water, cobalt, bromine and manganese within theliquid phase of the reaction medium as in Example 1. The operatingpressure of the reaction vessel overhead gas is again 0.52 megapascalgauge, and the operating temperature is again about 160° C. measurednear the mid-elevation of the reaction medium. Reaction mediumcomprising CTA is removed from the side of the reaction vessel at anelevation of 15 meters at a steady rate using an external de-aerationvessel.

The bubble column reaction vessel includes a vertical, cylindricalsection, with the inside diameter equal to 6.46 meters, keeping thesuperficial velocity of the gas phase approximately constant compared toExamples 1 and 2. In order to keep the same STR as Example 1, theoperating level is changed slightly to about 26.1 meters of reactionmedium. The height from lower TL to upper TL of the cylindrical sectionis 32 meters, the same as Example 1 and providing about the samefreeboard disengaging height between the top of the reaction medium andthe overhead gas outlet. The top and bottom of the cylindrical sectionare fitted with 2:1 elliptical heads. The height from the bottom of thereaction medium to the top of the cylindrical section is about 33.6meters, and the overall height of the reaction vessel is about 35.2meters. The feed is again distributed within the reaction vessel near anelevation of 2 meters above the bottom of the reaction medium using ahorizontal distributor assembly designed for a substantially uniformrelease of the feed over the area of cross section lying inside of aradius of 0.45*(inside diameter). The feed of air is again through anoxidant inlet distributor similar to the one shown in FIGS. 12-15, andall of the oxidant admission holes are located below the lower TL of thecylindrical section. The reflux solvent is distributed as droplets overessentially all of the cross sectional area of the disengaging zone.

The H:W ratio of the reaction medium is decreased markedly to 4.0. TheL:D ratio of the reaction vessel is decreased markedly to 5.2. Thevolume occupied by the reaction medium is about 828 cubic meters, andthe reaction vessel contains about 410,000 kilograms of slurry. Theratio of slurry mass to para-xylene feed rate is about 8.3 hours. TheSTR is about 59 kilograms of para-xylene fed per cubic meter of reactionmedium per hour. The pressure differential from the bottom of thereaction medium to the overhead off-gas exiting the reaction vessel isabout 0.13 megapascal. The upright surface area in contact with thereaction medium is about 546 square meters, which is about 3.24*Wmin*H.The ratio of the volume of reaction medium to upright surface area incontact with the reaction medium is about 1.52 meters. Under conditionsof this example, it is estimated that decomposition of acetic acid inthe reaction medium is desirably returned to the lower level of Example1, approximately 0.03 kilograms per kilogram of produced CTA.

However, the larger diameter of the reaction vessel employed in thisexample coupled with its smaller H:W ratio leads to very undesirableshifts in the flow velocities, mixing and staging within the reactionmedium. This leads to a significant increase in the loss of para-xylenein the overhead off-gas and in the formation of undesirable, coloredbyproducts. Simply stated, the axial flow velocities produced by naturalconvection forces grow larger in bubble columns of greater diameter evenwhen superficial velocity is kept constant. Calculated by a methodderived from Gas-Liquid Recirculation Model of Gupta and using aproprietary data base, the maximum upwards velocity, time-averaged, ofthe gas phase near the mid-elevation of the reaction medium is about 3.9meters per second. Similarly calculated, the maximum time-averagedupwards velocity of the slurry near the mid-elevation of the reactionmedium is about 2.2 meters per second. Similarly calculated, the maximumtime-averaged downwards velocity of the slurry in the outer annulus nearthe mid-elevation of the reaction medium is about 2.3 meters per second.

Since the height of the reaction medium is little changed between thisexample and Example 1, these increased time-averaged vertical velocitiescause the end-to-end mixing times to be significantly reduced in thisexample as compared to Example 1. This produces an undesirable increasein the amount of para-xylene migrating toward the top of the vesselbefore oxidizing. This leads to an undesirable loss in yield ofpara-xylene exiting the top of the reaction vessel with the off-gas, andit shifts more of the demand for dissolved molecular oxygen to nearerthe top of the reactor where the mole fraction of molecular oxygen isrelatively depleted within the gas phase. Furthermore, the increasedtime-averaged downward velocity of slurry in regions toward the vesselwall in this example causes more and larger bubbles of the gas phase tobe pulled downwards against their natural buoyancy in a gravitationalfield. This leads to an undesirable increase in the recirculation of gasphase partly depleted of molecular oxygen, which in turn leads toreduced availability of dissolved oxygen in these regions. Among othereffects, this reduced availability of dissolved oxygen in variousportions of the reaction medium leads to a significantly increasedformation ratio of undesirable, colored byproducts in this examplecompared to Example 1, and this elevated level of undesirable, coloredbyproducts renders the product unusable for many applications in PET.

Thus, Examples 2 and 3 demonstrate the inadequacy of the prior art fordesigning large scale oxidation bubble columns using primarilysuperficial gas velocity (Ug), L:D ratio, and average space-time rate ofreaction (STR).

Example 4

In this example, the pressure containing members of the bubble columnreaction vessel are the same as Example 3, but upright surfaces areadded within the reaction medium to impart vertical drag in order toestablish reaction staging profiles more similar to Example 1, therebyrestoring product quality and para-xylene yield, but without increasingthe decomposition ratio of acetic acid as in Example 2.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 15 meters at a steady rate using an externalde-aeration vessel.

The bubble column reaction vessel includes a vertical, cylindricalsection, with the inside diameter equal to 6.46 meters. The height fromlower TL to upper TL of the cylindrical section is 32 metrs, and theoperating level is about 26.3 meters of reaction medium. The top andbottom of the cylindrical section are fitted with 2:1 elliptical heads.The height from the bottom of the reaction medium to the top of thecylindrical section is about 33.6 meters, and the overall height of thereaction vessel is about 35.2 meters. The feed is again distributedwithin the reaction vessel near an elevation of 2 meters above thebottom of the reaction medium using a horizontal distributor assemblydesigned for a substantially uniform release of the feed over the areaof cross section lying inside of a radius of 0.45*(inside diameter). Thefeed of air is again through an oxidant inlet distributor similar to theone shown in FIGS. 12-15, and all of the oxidant admission holes arelocated below the lower TL of the cylindrical section. The refluxsolvent is distributed as droplets over essentially all of the crosssectional area of the disengaging zone.

The bubble column reaction vessel further includes two orthogonal planarsurfaces located with their line of intersection coincident with thevertical axis of symmetry of the cylindrical section. These planarsurfaces conveniently comprise plate metal of the same type and surfacefinish as is used in the cylindrical section of the reaction vessel.Each planar surface begins at a lower elevation of 3 meters above thelower TL and extends upwards by 20 meters. The two planar surfaces eachextend horizontally essentially all the way to the wall of thecylindrical section (i.e., the width of each is equal to insidediameter) and are supported from the cylindrical section. The thicknessand support of the planar surfaces are designed to withstand the variousforces that may occur in normal and upset operating conditions. Owing tothe volume occupied by the plate metal, the operating level is adjustedupwards slightly to 26.3 meters above the bottom of the reaction mediumin order to keep the same STR as Example 1.

Thus, in this example, the reaction medium is subdivided into 4 equallysized and shaped sub-volumes for 20 meters out of the total height ofthe reaction medium. These 4 sub-volumes communicate with each otherboth below and above the planar surfaces. Owing to the relativelyuniform distribution of oxidant and para-xylene feed below the lowerextremity of the planar surfaces, each of the 4 sub-volumes has asimilar superficial velocity of the gas phase and a similar reactionintensity profile. The 4 sub-volumes may be conceived as akin to 4smaller-sized bubble column reaction vessels within the shell of onepressure containing vessel.

The H:W ratio of the reaction medium is 4.1. The L:D ratio of thereaction vessel is 5.2. The volume occupied by the reaction medium isabout 828 cubic meters, and the reaction vessel contains about 410,000kilograms of slurry. The ratio of slurry mass to para-xylene feed rateis about 8.3 hours. The STR is about 59 kilograms of para-xylene fed percubic meter of reaction medium per hour. The pressure differential fromthe bottom of the reaction medium to the overhead off-gas exiting thereaction vessel is about 0.13 megapascal. The upright surface area incontact with the reaction medium is about 1,066-square meters, which isabout 6.29*Wmin*H. The ratio of the volume of reaction medium to uprightsurface area in contact with the reaction medium is about 0.78 meters.This value is intermediate between Examples 1 and 3, and it liesusefully closer to Example 1. Under conditions of this example, it isestimated that decomposition of acetic acid in the reaction medium isdesirably returned to the lower level of Example 1, approximately 0.03kilograms per kilogram of produced CTA.

The maximum time-averaged upwards and downwards velocities of the gasphase and of the slurry are reduced in this example as compared toExample 3. This provides useful improvements in the vertical profile ofpara-xylene, and it leads to improved availability of dissolved oxygenin the liquid phase near the upright wall surfaces. Together, thesechanges improve para-xylene yield and reduce the formation ofundesirable, colored byproducts in this example as compared to Example3.

Example 5

In this example, the pressure containing members of the reaction vesselare the same as Example 3, but non-fouling baffle members are addedwithin the reaction medium to re-establish reaction staging profilesmore similar to Example 1, thereby restoring product quality andpara-xylene yield, but without increasing the decomposition ratio ofacetic acid as in Example 2.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 15 meters at a steady rate using an externalde-aeration vessel.

The bubble column reaction vessel includes a vertical, cylindricalsection, with the inside diameter equal to 6.46 meters. The height fromlower TL to upper TL of the cylindrical section is 32 metrs, and theoperating level is about 26.1 meters of reaction medium. The top andbottom of the cylindrical section are fitted with 2:1 elliptical heads.The height from the bottom of the reaction medium to the top of thecylindrical section is about 33.6 meters, and the overall height of thereaction vessel is about 35.2 meters. The feed is again distributedwithin the reaction vessel near an elevation of 2 meters above thebottom of the reaction medium using a horizontal distributor assemblydesigned for a substantially uniform release of the feed over the areaof cross section lying inside of a radius of 0.45*(inside diameter). Thefeed of air is again through an oxidant inlet distributor similar to theone shown in FIGS. 12-15, and all of the oxidant admission holes arelocated below the lower TL of the cylindrical section. The refluxsolvent is distributed as droplets over essentially all of the crosssectional area of the disengaging zone.

The bubble column reaction vessel further includes a horizontal baffleassembly located in the bubble column at a height of 12 meters above thelower TL of the reaction vessel. This places the baffle assembly about13.6 meters, or 2.1*D, above the bottom of the reaction medium. Thisbaffle assembly is comprised of 15 individual baffle members. Eachbaffle member is comprised of an extruded or fabricated L-shape whereinboth legs of the L-shape are 0.15 meters wide and the included anglebetween the two legs is 90-degrees. The L-shapes are all arrangedhorizontal, parallel to each other, with the corners pointing upwards,and located at the same elevation. The two terminal edges of eachL-shape are all at the same elevation, below the upwards pointingcorners. When viewed on end, each member appears as an inverted V-shape.Thus, the percentage of the baffle assembly comprising upwardly-facingplanar surfaces inclined less than 5 degrees from horizontal iseffectively nil. The gap between the lower edges of each member and itsnear neighbor is always 0.21 meters. The longest of the members has alength effectively equal to the inside diameter of the cylindricalvessel, extending diametrically across the vessel from wall to wall. Theother 14 individual baffle members are all necessarily shorter inlength. All baffle members are supported on each end by extending allthe way to the cylindrical wall and attaching thereto. Thus, the openarea at the elevation of the baffle assembly is about 16-square meters,which is about 50-percent of the cross sectional area of the reactionvessel at that elevation. The baffle members are designed to withstandthe various forces that may occur in normal and upset operatingconditions. The members are constructed of the same metal as the pipingcomponents used in the air sparger assembly, which metal isappropriately selected to resist corrosion and erosion. However, thesurfaces of the baffle members are polished to a surface finish of 125RMS or finer. Despite the precipitation of about 76,000 kilograms perhour of CTA within the reaction vessel, the baffle assembly does notexcessively foul or slough away chunks of solids.

The H:W ratio of the reaction medium is 4.0. The L:D ratio of thereaction vessel is 5.2. The volume occupied by the reaction medium isabout 828 cubic meters, and the reaction vessel contains about 410,000kilograms of slurry. The ratio of slurry mass to para-xylene feed rateis about 8.3 hours. The STR is about 59 kilograms of para-xylene fed percubic meter of reaction medium per hour. The pressure differential fromthe bottom of the reaction medium to the overhead off-gas exiting thereaction vessel is about 0.13 megapascal. The upright surface area incontact with the reaction medium is about 546 square meters, which isabout 3.24*Wmin*H. The ratio of the volume of reaction medium to uprightsurface area in contact with the reaction medium is about 1.52 meters.Under conditions of this example, it is estimated that decomposition ofacetic acid in the reaction medium is desirably returned to the lowerlevel of Example 1, approximately 0.03 kilograms per kilogram ofproduced CTA.

The effect of the horizontal baffle assembly is to disrupt the verticalvelocity of the gas phase and of the slurry within the reaction vessel.This retards the progress of para-xylene toward the top surface of thereaction medium, leading to a beneficial reduction in yield loss ofpara-xylene in the overhead off-gas. Additionally, the staging ofmolecular oxygen and oxidizable compound are improved providing areduction in the formation of undesirable, colored byproducts in thisexample as compared to Example 3.

Example 6

In this example, the reaction vessel is designed for very high gassuperficial velocities and gas hold-up values according to the presentinvention. Using a smaller D enables a higher L:D ratio withoutresorting to an excessively tall reaction vessel and without incurringexcessive decomposition of acetic acid solvent.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 28 meters at a steady rate using an externalde-aeration vessel.

The bubble column reaction vessel includes a vertical, cylindricalsection, with the inside diameter equal to 5.00 meters. The height fromlower TL to upper TL of the cylindrical section is 70 meters. The topand bottom of the cylindrical section are fitted with 2:1 ellipticalheads. The height from the bottom of the reaction medium to the top ofthe cylindrical section is about 71.3 meters, and the overall height ofthe reaction vessel is about 72.5 meters. The operating level is about61.3 meters above the bottom of the reaction medium. The feed is againdistributed within the reaction vessel near an elevation of 2 metersabove the bottom of the reaction medium using a horizontal distributorassembly designed for a substantially uniform release of the feed overthe area of cross section lying inside of a radius of 0.45*(insidediameter). The feed of air is again through an oxidant inlet distributorsimilar to the one shown in FIGS. 12-15, and all of the oxidantadmission holes are located below the lower TL of the cylindricalsection. The reflux solvent is distributed as droplets over essentiallyall of the cross sectional area of the disengaging zone.

The H:W ratio of the reaction medium is 12.3. The L:D ratio of thereaction vessel is 14.3. The volume occupied by the reaction medium isabout 1,190 cubic meters, and the reaction vessel contains about 420,000kilograms of slurry. The ratio of slurry mass to para-xylene feed rateis about 8.7 hours. The STR is about 41 kilograms of para-xylene fed percubic meter of reaction medium per hour. The pressure differential fromthe bottom of the reaction medium to the overhead off-gas exiting thereaction vessel is about 0.21 megapascal. The upright surface area incontact with the reaction medium is about 975 square meters, which isabout 3.18*Wmin*H. The ratio of the volume of reaction medium to uprightsurface area in contact with the reaction medium is about 1.22 meters.The relatively small value of D produces a superficial velocity of thegas phase at the mid-height of the reaction medium that is about 1.7times the superficial velocity used in Examples 1 through 5. The gashold-up at the mid-elevation of the reaction medium is in excess of 0.6.Under conditions of Example 6, it is estimated that decomposition ofacetic acid in the reaction medium is desirably reduced below 0.03kilograms per kilogram of produced CTA. This is owing to the reducedamount of slurry, more specifically acetic acid, in the reaction vesselcompared to Example 3.

In this example, the H:W ratio is favorable for reduced end-to-endmixing and for beneficial staging of molecular oxygen and oxidizablecompound. However, the axial velocities are higher than Example 1,accelerating end-to-end mixing for a given H:W. Fortunately, the lowerSTR reduces the volumetric demand for molecular oxygen transfer from gasto liquid; and the increased gas hold-up serves to increase thecapability to transfer molecular oxygen from gas to liquid. On balance,the level of production of undesirable, colored byproducts is estimatedto be comparable to Example 1.

Example 7

In this example, the reaction vessel is designed for yet higher gassuperficial velocities and gas hold-up values according to the presentinvention. An enlarged, surmounting disengaging zone is used to limitentrainment of slurry in overhead off-gas.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 28 meters at a steady rate using an externalde-aeration vessel.

The bubble column reaction vessel includes a vertical, cylindricalsection, with the inside diameter equal to 4.60 meters. The height fromlower TL to the upper end of the cylindrical section is 60 meters. Atthe upper end of this cylindrical section, a conical section diverges toan inside diameter of 7 meters while rising in height by 2 meters. Theslope of the conical wall is thus about 31 degrees from vertical.Surmounting the conical section is a gas-disengaging cylindrical sectionwith an inside diameter of 7 meters. The height of the upper cylindricalsection is 7 meters. The vessel is fitted with 2:1 elliptical heads atthe top and bottom. Thus, the combined height of the reaction vessel isabout 71.9 meters. The operating level is about 61.2 meters above thebottom of the reaction medium, placing it near the conjunction of themain cylindrical body and the diverging conical section. The feed isagain distributed within the reaction vessel near an elevation of 2meters above the bottom of the reaction medium using a horizontaldistributor assembly designed for a substantially uniform release of thefeed over the area of cross section lying inside of a radius of0.45*(inside diameter). The feed of air is again through an oxidantinlet distributor similar to the one shown in FIGS. 12-15, and all ofthe oxidant admission holes are located below the lower TL of the lowercylindrical section. The reflux solvent is distributed as droplets overessentially all of the cross sectional area of the expanded disengagingsection.

The H:W ratio of the reaction medium and the L:D ratio of the reactionvessel are 13.3. The X:D ratio of the reaction vessel is 1.5. The L:Yratio of the reaction vessel is 5.7. The volume occupied by the reactionmedium is about 1,000 cubic meters, and the reaction vessel containsabout 320,000 kilograms of slurry. The ratio of slurry mass topara-xylene feed rate is about 6.5 hours. The STR is about 49 kilogramsof para-xylene fed per cubic meter of reaction medium per hour. Thepressure differential from the bottom of the reaction medium to theoverhead off-gas exiting the reaction vessel is about 0.19 megapascal.The upright surface area in contact with the reaction medium is about896 square meters, which is about 3.19*Wmin*H. The ratio of the volumeof reaction medium to upright surface area in contact with the reactionmedium is about 1.12 meters. The relatively small value of D produces asuperficial velocity of the gas phase at the mid-height of the reactionmedium that is about 2 times the superficial velocity used in Examples 1through 5. However, the superficial velocity of the gas phase in theexpanded disengaging section is reduced to about 0.85 times thesuperficial velocity used in Examples 1 through 5. The gas hold-up atthe mid-elevation of the reaction medium is about 0.7. Under conditionsof this example, it is estimated that decomposition of acetic acid inthe reaction medium is desirably reduced below 0.03 kilograms perkilogram of produced CTA. This is owing to the reduced amount of slurry,more specifically acetic acid, in the reaction vessel compared toExample 3. It is estimated that the level of undesirable coloredbyproducts is lower than in Example 6 owing to the improved staging andhigher gas hold-up.

Example 8

In this example, the reaction vessel is the same as Example 7, but theoperating level is raised to be about 63.2 meters above the bottom ofthe reaction medium, placing it near the conjunction of the divergingconical section and the expanded gas-disengaging cylindrical section.This provides various advantages versus controlling the level in themain cylindrical body section, including a reduced tendency for the topof the reaction medium to be foamy and too poorly circulating.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 28 meters at a steady rate using an externalde-aeration vessel.

The H:W ratio of the reaction medium is about 13.7, and the L:D ratio ofthe reaction vessel is 13.3. The volume occupied by the reaction mediumis about 1,060 cubic meters, and the reaction vessel contains about330,000 kilograms of slurry. The ratio of slurry mass to para-xylenefeed rate is about 6.8 hours. The STR is about 46 kilograms ofpara-xylene fed per cubic meter of reaction medium per hour. Thepressure differential from the bottom of the reaction medium to theoverhead off-gas exiting the reaction vessel is about 0.20 megapascal.The upright surface area in contact with the reaction medium is about953 square meters, which is about 3.39*Wmin*H. The ratio of the volumeof reaction medium to upright surface area in contact with the reactionmedium is about 1.11 meters. The superficial velocity of the gas phaseat the mid-height of the reaction medium is about 2 times thesuperficial velocity used in Examples 1 through 5. The gas hold-up atthe mid-elevation of the reaction medium is about 0.7. Under conditionsof this example, it is estimated that decomposition of acetic acid inthe reaction medium is desirably reduced below 0.03 kilograms perkilogram of produced CTA. This is owing to the reduced amount of slurry,more specifically acetic acid, in the reaction vessel compared toExample 3. It is estimated that the level of undesirable coloredbyproducts is lower than in Example 6 owing to the improved staging andgas hold-up.

Example 9

In this example, the pressure containing members of the reaction vesselare the same as Example 7, but the internal members used to introducethe oxidant and the para-xylene are importantly modified to providemultiple vertically separated entries for each.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 28 meters at a steady rate using an externalde-aeration vessel.

The bubble column reaction vessel includes the same oxidant distributorin the bottom head of the reaction vessel as in Example 8. However, only70 percent of the total gas-phase oxidant stream is introduced via thislower distributor. The other 30 percent of the gas-phase oxidant isintroduced via an elevated oxidant inlet distributor. This flow ratio isimposed by flow control loops using control valves and flow transmittersconveniently located on the supply conduits for the compressed airexternal to the reaction vessel. The elevated oxidant distributorcomprises a horizontal, mitered, square-shaped flow conduit, rather thanan octagonal one as is used in the lower elliptical head. Thesquare-shaped conduit conveniently comprises nominal 14-inch Schedule10S piping materials. The distance from the centroid of one side to thecentroid of the opposite side is 1 meter. The elevated oxidantdistributor comprises about sixty release holes for gas-phase oxidant,all 0.03 meters in diameter and near the bottom of the conduit about 14meters above the bottom of the reaction medium. This elevated oxidantdistributor serves at least two useful functions. Firstly, the oxidantflow injected downwards into the reaction medium disrupts the axialvelocity profile rising along the vertical axis of symmetry of thecylindrical section. This imposes a useful hydraulic baffling to slowthe spread of para-xylene into upper regions of the reaction medium,relating to overhead yield loss and to reducing demand for dissolvedoxygen in upper regions. A few meters above the elevated oxidant inlet,the natural convention flow pattern re-organizes itself to race upwardsalong the central axis of symmetry, but the hydraulic baffling isnonetheless efficacious. Secondly, most of the heat of reaction isremoved from the reaction medium via the evaporation of solvent, andmost of this evaporation occurs near the oxidant feed locations. Byseparating vertically the location of introduction of parts of thegas-phase oxidant stream, the vertical profile of temperature in thereaction medium is adjusted.

The bubble column reaction vessel includes two para-xylene inletdistributors similar to the one in Example 8. The lower para-xyleneinlet distributor is located to provide a substantially uniform releaseof 50 percent of the liquid-phase feed over the area of cross sectionlying inside of a radius of 0.45*(inside diameter) at a lower elevationof 2 meters above the bottom of the reaction medium. The upperpara-xylene inlet distributor is located to provide a substantiallyuniform release of 50 percent of the liquid-phase feed over the area ofcross section lying inside of a radius of 0.45*(inside diameter) athigher elevation of 10 meters above the bottom of the reaction medium.This flow ratio is imposed by flow control loops using control valvesand flow transmitters conveniently located on the supply conduits forthe liquid-phase feed external to the reaction vessel.

In this example, the operating level is raised to be about 63.7 metersabove the bottom of the reaction medium, placing it just above thediverging conical section and into the expanded gas-disengagingcylindrical section. The H:W ratio of the reaction medium is about 13.8,and the L:D ratio of the reaction vessel is 13.3. The volume occupied bythe reaction medium is about 1,070 cubic meters, and the reaction vesselcontains about 340,000 kilograms of slurry. The ratio of slurry mass topara-xylene feed rate is about 6.9 hours. The STR is about 46 kilogramsof para-xylene fed per cubic meter of reaction medium per hour. Thepressure differential from the bottom of the reaction medium to theoverhead off-gas exiting the reaction vessel is about 0.20 megapascal.The upright surface area in contact with the reaction medium is about975 square meters, which is about 3.47*Wmin*H. The ratio of the volumeof reaction medium to upright surface area in contact with the reactionmedium is about 1.10 meters. The superficial velocity of the gas phaseat the mid-height of the reaction medium is about 2 times thesuperficial velocity used in Examples 1 through 5.

Example 10

In this example, the reaction vessel is designed with three differentcylindrical diameters at different elevations, with the mid-elevationdiameter being smallest. This configuration benefits the lowercylindrical section, where the liquid feed stream and gas-phase oxidantfirst enter, with a relatively larger mass of liquid phase for initialdilution and reaction of para-xylene where oxygen is still moreabundant; the middle cylindrical section, where molecular oxygen isincreasingly depleted, with a relatively greater gas hold-up andgas-to-liquid mass transfer rate; and the top cylindrical section, whichis a gas-disengaging zone, with relatively reduced gas-phase velocity tolimit entrainment of slurry in overhead off-gas.

The feed rate of para-xylene is again 49,000 kilograms per hour—7 timeslarger than in Example 1. The other feed flows are increased with thesame 7:1 ratio to Example 1. The compositions of the feeds are the sameas in Example 1, providing the same concentrations of water, cobalt,bromine and manganese within the liquid phase of the reaction medium asin Example 1. The operating pressure of the reaction vessel overhead gasis again 0.52 megapascal gauge, and the operating temperature is againabout 160° C. measured near the mid-elevation of the reaction medium.Reaction medium comprising CTA is removed from the side of the reactionvessel at an elevation of 28 meters at a steady rate using an externalde-aeration vessel.

The bubble column includes three vertical, cylindrical sections ofdifferent diameters. The lowest cylindrical section has an insidediameter of 6.46 meters, giving a superficial velocity of the gas phasein this section approximately equal to the superficial velocity ofExample 1. The height of this lower cylindrical section from lower TL tothe upper end is 8 meters. At the upper end of this lower cylindricalsection, a conical section converges to an inside diameter of 4.5 meterswhile rising in height by 1 meter. The slope of this conical wall isthus about 44 degrees from vertical. Surmounting the lower conicalsection, a middle cylindrical section has an inside diameter of 4.5meters, giving a superficial velocity of the gas phase in this sectionabout twice the superficial velocity in the lowest cylindrical section.The height of the middle cylindrical section is 45 meters. At the upperend of the middle cylindrical section, a conical section diverges to aninside diameter of 7 meters while rising in height by 2 meters. Theslope of the conical wall is thus about 32 degrees from vertical.Surmounting the upper conical diverging section is a gas-disengagingcylindrical section with an inside diameter of 7 meters. The height ofthe upper cylindrical section is 7 meters. The vessel is fitted with 2:1elliptical heads at the top and bottom. Thus, the combined height of thereaction vessel is about 66.4 meters. The operating level is about 57.6meters above the bottom of the reaction medium, placing it near theconjunction of the diverging conical section and the upper cylindricalsection. The feed is again distributed within the reaction vessel nearan elevation of 2 meters above the bottom of the reaction medium using ahorizontal distributor assembly designed for a substantially uniformrelease of the feed over the area of cross section lying inside of aradius of 0.45*(inside diameter). The feed of air is again through anoxidant inlet distributor similar to the one shown in FIGS. 12-15, andall of the oxidant admission holes are located below the lower TL of thelowest cylindrical section. The reflux solvent is distributed asdroplets over essentially all of the cross sectional area of the topcylindrical section.

The volume occupied by the reaction medium is about 1,080 cubic meters,and the reaction vessel contains about 400,000 kilograms of slurry. Theratio of slurry mass to para-xylene feed rate is about 8.1 hours. TheSTR is about 45 kilograms of para-xylene fed per cubic meter of reactionmedium per hour. The pressure differential from the bottom of thereaction medium to the overhead off-gas exiting the reaction vessel isabout 0.20 megapascal. The upright surface area in contact with thereaction medium is about 944 square meters, which is about 3.64*Wmin*Hand about 2.34*Wmax*H. The ratio of the volume of reaction medium toupright surface area in contact with the reaction medium is about 1.14meters. The ratio of L_(l):D_(l) is about 1.5:1. The ratio ofL_(u):D_(u) is about 10:1. The ratio of L_(l):L_(u) is about 0.21:1. Theratio of X:D_(l) is about 1.1:1. The ratio of L_(u):Y is about 4.2:1.The ratio of L_(t):D_(l) is about 0.15:1.

The larger diameter at the base of the reactor provides a large slurrymass near the introduction zone for the feed para-xylene, where liquidflows and mixing are quite important to provide feed dilution to avoidcoupled, colored aromatic impurities. Also, this larger diameter placesa larger fraction of the reaction medium under more head pressure fromthe slurry above, promoting oxygen partial pressure and mass transfer ofmolecular oxygen from gas to liquid. The smaller diameter, elongatedmiddle cylindrical section provides reactant staging and high gashold-up; this improves the match of reaction demand for dissolved oxygenwith the mass transfer supply from the rising gas phase, wherein oxygenis increasingly consumed and its partial pressure declining. Theinvention has been described in detail with particular reference topreferred embodiments thereof, but will be understood that variationsand modification can be effected within the spirit and scope of theinvention.

1. A process comprising: oxidizing an oxidizable compound in a liquidphase of a reaction medium contained in an agitated reactor, whereinsaid reaction medium has a gas hold-up of at least about 0.6 on atime-averaged and volume-averaged basis.
 2. The process of claim 1wherein said reaction medium comprises at least about 3 weight percentsolids on a time-averaged and volume-averaged basis.
 3. The process ofclaim 1 wherein said reactor is a bubble column reactor, wherein saidreaction medium has a gas hold-up in the range of from 0.65 to 0.85 on atime-averaged and volume-averaged basis, wherein said reaction mediumcomprises in the range of from about 5 to about 40 weight percent solidson a time-averaged and volume-averaged basis.
 4. The process of claim 1wherein the time-averaged superficial velocity of a gas phase of saidreaction medium at the half height of said reaction medium is in therange of from about 0.8 to about 5 meters per second.
 5. The process ofclaim 1 wherein said oxidizable compound is an aromatic compound.
 6. Theprocess of claim 1 wherein said oxidizable compound is selected from thegroup consisting of para-xylene, meta-xylene, para-tolualdehyde,meta-tolualdehyde, para-toluic acid, meta-toluic acid, acetaldehyde, andcombinations of two or more thereof.
 7. The process of claim 1 whereinsaid oxidizable compound is para-xylene.
 8. The process of claim 1wherein said reaction medium is a three-phase reaction medium.
 9. Theprocess of claim 1 wherein said oxidizing causes the formation of solidsin said reaction medium.
 10. The process of claim 1 wherein saidoxidizing causes at least about 10 weight percent of said oxidizablecompound to form solids in said reaction medium.
 11. The process ofclaim 1 wherein said reactor is a bubble column reactor.
 12. The processof claim 11 wherein said oxidizable compound comprises para-xylene. 13.The process of claim 1 wherein said oxidizing is carried out in thepresence of a catalyst system comprising cobalt.
 14. The process ofclaim 13 wherein said catalyst system further comprises bromine andmanganese.
 15. The process of claim 1 wherein said reaction medium has amaximum height (H), a maximum width (W), and an H:W ratio of at leastabout 3:1.
 16. The process of claim 15 wherein said H:W ratio is in therange of from about 8:1 to about 20:1.
 17. The process of claim 15wherein said process further comprises introducing a predominatelygas-phase oxidant stream comprising molecular oxygen into said reactionzone.
 18. The process of claim 17 wherein said oxidant stream comprisesless than about 50 mole percent molecular oxygen.
 19. The process ofclaim 17 wherein a majority of said molecular oxygen enters saidreaction zone within about 0.25 W of the bottom of said reaction zone.20. The process of claim 17 wherein a majority of said molecular oxygenenters said reaction zone within about 0.025 H of the bottom of saidreaction zone.
 21. The process of claim 17 wherein said process furthercomprises introducing a predominately liquid-phase feed streamcomprising said oxidizable compound into said reaction zone.
 22. Theprocess of claim 21 wherein said reactor is a bubble column reactor. 23.The process of claim 21 wherein at least about 50 weight percent of saidoxidizable compound enters said reaction zone within about 2.5 W of thelowest location where said molecular oxygen enters said reaction zone.24. The process of claim 21 wherein said feed stream is introduced intosaid reaction zone via a plurality of feed openings, wherein at leasttwo of said feed openings are vertically spaced from one another by atleast about 0.5 W.
 25. The process of claim 24 wherein at least two ofsaid feed openings are vertically-spaced from one another by a distanceof at least about 1.5 W.
 26. The process of claim 21 wherein said feedstream is introduced into said reaction zone in a manner such that whensaid reaction zone is theoretically partitioned into 4 verticalquadrants of equal volume by a pair of intersecting vertical planes, notmore than about 80 weight percent of said oxidizable compound enterssaid reaction zone in a single one of said vertical quadrants.
 27. Theprocess of claim 26 wherein not more than about 60 weight percent ofsaid oxidizable compound enters said reaction zone in a single one ofsaid vertical quadrants.
 28. The process of claim 1 wherein at least aportion of said reaction zone is defined by one or more uprightsidewalls of said reactor, wherein at least about 25 weight percent ofsaid oxidizable compound enters said reaction zone at one or morelocations spaced inwardly at least 0.05 D from said upright sidewalls,wherein said reaction zone has a maximum diameter (D).
 29. The processof claim 28 wherein at least about 50 weight percent of said oxidizablecompound enters said reaction zone at one or more locations spacedinwardly at least 0.05 D from said upright sidewalls.
 30. The process ofclaim 28 wherein said reactor is a bubble column reactor.
 31. A processfor producing terephthalic acid, said process comprising: (a) oxidizingpara-xylene a liquid phase of a three-phase reaction medium containedwithin a bubble column reactor to thereby form crude terephthalic acid,wherein said reaction medium has a gas hold-up of at least about 0.6 ona time-averaged and volume-averaged basis; and (b) oxidizing at least aportion of said crude terephthalic acid in a secondary oxidation reactorto thereby form purer terephthalic acid.
 32. The process of claim 31wherein said reaction medium has a gas hold-up in the range of from 0.65to 0.85 on a time-averaged and volume-averaged basis.
 33. The process ofclaim 31 wherein said oxidizing in said secondary oxidation reactordecreases the average concentration of 4-CBA present in said crudeterephthalic acid by at least about 200 ppmw to thereby form said purerterephthalic acid.
 34. The process of claim 31 wherein said crudeterephthalic acid has an average 4-CBA concentration of at least about400 ppmw and said purer terephthalic acid has an average 4-CBAconcentration of less than about 400 ppmw.
 35. The process of claim 31wherein said oxidizing in said secondary oxidation reactor decreases theaverage concentration of 4-CBA in said crude terephthalic acid by atleast about 400 ppmw to thereby form said purer terephthalic acid,wherein said crude terephthalic acid has an average 4-CBA concentrationof at least about 800 ppmw and said purer terephthalic acid has anaverage 4-CBA concentration of less than about 250 ppmw.
 36. The processof claim 31 wherein said oxidizing in said secondary oxidation reactoris carried out at an average temperature at least about 110° C. greaterthan the average temperature of said oxidizing in said bubble columnreactor.
 37. The process of claim 36 wherein said oxidizing in saidbubble column reactor is carried out at an average temperature in therange of from about 125 to about 200° C.
 38. The process of claim 36wherein said oxidizing in said secondary oxidation reactor is carriedout at an average temperature in the range of from about 160 to about240° C.
 39. The process of claim 31 wherein said oxidizing in saidsecondary oxidation reactor is carried out at an average temperature inthe range of from about 20 to about 80° C. greater than the averagetemperature of said oxidizing in said bubble column reactor, whereinsaid oxidizing in said bubble column reactor is carried out at anaverage temperature in the range of from about 140 to about 180° C.,wherein said oxidizing in said secondary oxidation reactor is carriedout at an average temperature in the range of from about 180 to about220° C.
 40. The process of claim 31 wherein a substantial portion ofsaid crude terephthalic acid exists as solid crude terephthalic acidparticles having an average BET surface area of at least about 0.6meters squared per gram.
 41. The process of claim 40 wherein said solidcrude terephthalic acid particles have an average particle size in therange of from about 20 to about 150 microns.
 42. The process of claim 41wherein a substantial portion of said solid crude terephthalic acidparticles are formed of a plurality of agglomerated sub-particles havingan average particle size in the range of from about 0.5 to about 30microns.
 43. The process of claim 42 wherein said solid crudeterephthalic acid particles have an average particle size in the rangeof from about 30 to about 120 microns, wherein said sub-particles havean average particle size in the range of from about 1 to about 15microns.
 44. The process of claim 31 wherein said process furthercomprises recovering an initial slurry comprising a mother liquor andsaid crude terephthalic acid from said bubble column reactor, whereinsaid process further comprises replacing at least about 50 weightpercent of said mother liquor in said initial slurry with a replacementsolvent to thereby provide a solvent-exchanged slurry comprising saidreplacement solvent and said crude terephthalic acid, wherein saidprocess further comprises introducing said solvent-exchanged slurry intosaid secondary oxidation reactor.
 45. An oxidation product produced bythe process of claim
 1. 46. A terephthalic acid product produced by theprocess of claim 31.